SEPARATION OF CARBON DIOXIDE (CO2) FROM GAS MIXTURES BY CALCIUM BASED REACTION SEPARATION (CaRS-CO2) PROCESS

ABSTRACT

A reaction-based process developed for the selective removal of CO2 from a multicomponent gas mixture to provide a gaseous stream depleted in CO2 compared to the inlet CO2 concentration. The proposed process effects the separation of CO2 from a mixture of gases by its reaction with metal oxides. The Calcium based Reaction Separation for CO2 (CaRS—CO2) process consists of contacting CO2 laden gas with CaO in a reactor such that CaO captures CO2 by the formation of CaCO3. CaCO3 is regenerated by calcination leading to the formation of fresh CaO sorbent and the evolution of a concentrated stream of CO2. The “regenerated” CaO is then recycled for the further capture of CO2. This carbonation-calcination cycle forms the basis of the CaRS—CO2 process. This process also may use a mesoporous CaCO3 structure that attains &gt;90% conversion over multiple carbonation and calcination cycles.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a continuation of U.S. patent application Ser. No.13/555,054 filed Jul. 20, 2012, which is a divisional of U.S. patentapplication Ser. No. 11/994,316 filed Apr. 14, 2008 and now U.S. Pat.No. 8,226,917, which is a 371 national stage filing of PCT/US2006/025266filed Jun. 28, 2006. International Patent Application No.PCT/US2006/025266 filed Jun. 28, 2006 is a continuation of U.S. patentapplication Ser. No. 11/255,099 filed Oct. 20, 2005 and now U.S. Pat.No. 7,618,606, which is a continuation-in-part of U.S. patentapplication Ser. No. 10/359,763 filed Feb. 6, 2003 now U.S. Pat. No.7,067,456. International Patent Application No. PCT/US2006/025266 filedJun. 28, 2006 also claims benefit of U.S. Patent Application No.60/694,594 filed Jun. 28, 2005. All aforementioned applications arehereby incorporated herein by reference.

TECHNICAL FIELD OF THE INVENTION

The present invention relates to the application of chemical sorbentsfor the separation of CO₂ from gas mixtures.

BACKGROUND OF THE INVENTION

As used herein, the term “supersorbent” shall mean a sorbent as taughtin U.S. Pat. No. 5,779,464 entitled “Calcium Carbonate Sorbent andMethods of Making and Using Same”, the teachings of which are herebyincorporated by reference.

As used herein, the term “microporous” shall mean a pore sizedistribution of less than 5 nanometers. As used herein, the term“mesoporous” shall mean a pore size distribution of from about 5nanometers to about 20 nanometers.

Atmospheric CO₂ concentration has been increasing steadily since theindustrial revolution. It has been widely accepted that the while theCO₂ concentration was about 280 ppm before the industrial revolution, ithas increased from 315 ppmv in 1959 to 370 ppmv in 2001 [Keeling, C. D.and T. P. Whorf. 2002. Atmospheric CO₂ records from sites in the SIO airsampling network. In Trends: A Compendium of Data on Global Change.Carbon Dioxide Information Analysis Center, Oak Ridge NationalLaboratory, U.S. Department of Energy, Oak Ridge, Tenn., U.S.A. Thisdata is also available fromhttp://cdiac.esd.ornl.gov/ftp/maunaloa-co2/maunaloa.co2]. Rising CO₂concentrations has been reported to account for half of the greenhouseeffect that causes global warming [IPCC Working Group I. IPCC ClimateChange 1995—The Science of Climate Change: The Second Assessment Reportof the Intergovernmental Panel on Climate Change; Houghton, J. T., MeiraFilho, L. G., Callander, B. A., Harris, N., Kattenberg, A., Maskell K,Eds.; Cambridge University Press: Cambridge, U.K., 1996]. Although theanthropogenic CO₂ emissions are small compared to the amount of CO₂exchanged in the natural cycles, the discrepancy between the long lifeof CO₂ in the atmosphere (50-200 years) and the slow rate of natural CO₂sequestration processes leads to CO₂ build up in the atmosphere. TheIPCC (Intergovernmental Panel on Climate Change) opines that “thebalance of evidence suggests a discernible human influence on the globalclimate.” Therefore, it is necessary to develop cost effective CO₂management schemes to curb its emission.

Many of the envisaged CO₂ management schemes consist of threeparts—separation, transportation and sequestration of CO₂ [FETC CarbonSequestration R&D Program Plan: FY 1999-2000. National Energy TechnologyLaboratory, Department of Energy, Washington, D.C., 1999]. The cost ofseparation and compression of CO₂ to 110 bar (for transportation of CO₂in liquid state) is estimated at $30-50 per ton CO₂, and transportationand sequestration would cost about $1-3 per ton per 100 km and $1-3 perton of CO₂, respectively [Wallace, D. Capture and Storage of CO₂. WhatNeeds To Be Done. Presented at the 6th Conference of the Parties, COP 6,to the United Nations Framework Convention on Climate Change; The Hague,The Netherlands, Nov. 13-24, 2000; www.iea.org/envissu/index.htm]. Thecapture of CO₂ imposes severe energy penalties thereby reducing the netelectricity output by as much as 13-37% [Herzog, H.; Drake, E.; Adams,E. CO₂ Capture, Reuse, and Storage Technologies for Mitigating GlobalClimate Change. A White Paper; Final Report No. DE-AF22-96PC01257,January 1997]. The dominating costs associated with the current CO₂separation technologies necessitate development of economicalalternatives.

Historically, CO₂ separation was motivated by enhanced oil recovery[Kaplan, L. J. Cost-Saving Processes Recovers CO₂ from Power-Plant Fluegas. Chem. Eng. 1982, 89 (24), 30-31; Pauley, C. P.; Smiskey, P. L.;Haigh, S. N—ReN Recovers CO₂ from Flue Gas Economically. Oil Gas J.1984, 82(20), 87-92]. Currently, industrial processes such as limestonecalcination, synthesis of ammonia and hydrogen production require CO₂separation. Absorption processes employ physical and chemical solventssuch as Selexol and Rectisol, MEA and KS-2 [Reimer, P.; Audus, H.;Smith, A. Carbon Dioxide Capture from Power Stations. IEA Greenhouse R&DProgramme, www.ieagreen.org.uk, 2001. ISBN 1 898373 15 9; Blauwhoff, P.M. M.; Versteeg, G. F.; van Swaaij, W. P. M. A study on the reactionbetween CO₂ and alkanoamines in aqueous solution. Chem. Eng. Sci. 1984,39(2), 207-225. Mimura, T.; Simayoshi, H.; Suda, T.; Iijima, M.;Mitsuake, S. Development of Energy Saving Technology for Flue Gas CarbonDioxide Recovery by Chemical Absorption Method and Steam System in PowerPlant. Energy Convers. Mgmt. 1997, 38, Suppl. P. S57-S62]. Adsorptionsystems capture CO₂ on a bed of adsorbent materials such as molecularsieves and activated carbon [Kikkinides, E. S.; Yang, R. T.; Cho, S. H.Concentration and Recovery of CO₂ from flue gas by pressure swingadsorption. Ind. Eng. Chem. Res. 1993, 32, 2714-2720]. CO₂ can also beseparated from the other gases by condensing it out at cryogenictemperatures. Polymers, metals such as palladium, and molecular sievesare being evaluated for membrane based separation processes [Reimer, P.;Audus, H.; Smith, A. Carbon Dioxide Capture from Power Stations. IEAGreenhouse R&D Programme, www.ieagreen.org.uk, 2001. ISBN 1 898373 159].

Reaction based processes, as promulgated in this work, can be applied toseparate CO₂ from gas mixtures. This process is based on a heterogeneousgas-solid non-catalytic carbonation reaction where gaseous CO₂ reactswith solid metal oxide (represented by MO) to yield the metal carbonate(MCO₃). The reaction can be represented by:

MO+CO₂→MCO₃  (1)

Once the metal oxide has reached its ultimate conversion, it can bethermally regenerated to the metal oxide and CO₂ by the calcination ofthe metal carbonate product. The calcination reaction can be representedby:

MCO₃→MO+CO₂  (2)

As an example of the above-mentioned scheme, FIG. 1 shows the variationin the free energy of the carbonation reaction as a function oftemperature for calcium oxide. From the figure, we can see that thecarbonation reaction is thermodynamically favored with a decrease intemperature (Gibbs free energy declines with a decrease in temperature).However, at lower temperatures, the carbonation reaction is kineticallyslow. In fact, it takes geological time scales for the formation ofCaCO₃ by the reaction between CaO and atmospheric CO₂ (at 280-360 ppm)at ambient temperatures. It should also be noted that the carbonationreaction would be favored as long as the free energy is negative. Thiscreates an upper bound of 890° C. for carbonation to occur under a CO₂partial pressure of 1 atm. The equilibrium temperature for this reactionis a function of the partial pressure of CO₂. A reaction based CO₂separation process offers many advantages. Under ideal conditions, MEAcaptures 60 g CO₂/kg, silica gel adsorbs 13.2 g CO₂/kg and activatedcarbon adsorbs 88 g CO₂/kg. The sorption capacity of some metal oxides(such as the modified CaO, presented in this study) is about 700 gCO₂/kg of CaO. This is about an order of magnitude higher than thecapacity of adsorbents/solvents used in other CO₂ separation processesand would significantly reduce the size of the reactors and the materialhandling associated with CO₂ separation.

Numerous metal oxides exhibit the carbonation and calcination reaction.The calcination temperature of a few metal carbonates (CaCO₃˜750° C.,MgCO₃˜385° C., ZnCO₃˜340° C., PbCO₃˜350° C., CuCO₃˜225-290° C. andMnCO₃˜440° C.) makes them viable candidates for this process. Apart fromCaO, gas-solid carbonation of other metal oxides has not been widelystudied. The carbonation of ZnO to ZnCO₃ at 8-13° C. was low whenexposed to CO₂ and H₂O for over 100 days (Sawada, Y.; Murakami, M.;Nishide, T. Thermal analysis of basic zinc carbonate. Part 1.Carbonation process of zinc oxide powders at 8 and 13° C. Thermochim.Acta. 1996, 273, 95-102.). MnCO₃ undergoes a more complex thermaldegradation phenomena. MnCO₃ first decomposes to MnO₂ at 300° C., whichin turn changes to Mn₂O₃ at 440° C. At higher temperatures (˜900° C.),the final thermal decomposition product was identified as Mn₃O₄(Shaheen, W. M.; Selim, M. M. Effect of thermal treatment onphysicochemical properties of pure and mixed manganese carbonate andbasic copper carbonate. Thermochim. Acta. 1998, 322(2), 117-128.).Different oxides of manganese provide the flexibility of exploiting thecarbonation/calcination reaction over a wider temperature range. Aqueousphase MgO carbonation has been studied for its suitability formineral-based CO₂ sequestration (Fernandez, A. I.; Chimenos, J. M.;Segarra, M.; Fernandez, M. A.; Espiell, F. Kinetic study of carbonationof MgO slurries. Hydrometallurgy. 1999, 53, 155-167). The carbonationextent of Mg(OH)₂ was about 10% between 387-400° C. and 6% formationbetween 475-500° C. (Butt, D. P.; Lackner, K. S.; Wendt, C. H.; Conzone,S. D.; Kung, H.; Lu, Y-C.; Bremser, J. K. Kinetics of ThermalDehydroxylation and Carbonation of Magnesium Hydroxide. J. Am. Ceram.Soc. 1996, 79(7), 1892-1898). They attributed the low conversions to theformation of a non-porous carbonate product layer. This layer hindersthe inward diffusion of CO₂ and the outward diffusion of H₂O (a productof the carbonation reaction) leading to low conversions. The carbonationof PbO was studied as a part of the chemical heat pump process (Kato,Y.; Saku, D.; Harada, N.; Yoshizawa, Y. Utilization of High TemperatureHeat from Nuclear Reactor using Inorganic Chemical Heat Pump. Progressin Nuclear Energy. 1998, 32(3-4), 563-570. & Kato, Y.; Harada, N.;Yoshizawa, Y. Kinetic feasibility of a chemical heat pump for heatutilization from high temperature processes. Applied ThermalEngineering. 1999, 19, 239-254). They reported 30% conversion in an hourunder 100% CO₂ atmosphere at 300° C. Furthermore, they found thereactivity of PbO to drop with the number of carbonation-calcinationcycles.

Carbonation of calcium oxide has been widely studied. Relatedapplications of the CaO carbonation and calcination include the storageof energy (Barker, R. The Reversibility of the Reaction CaCO₃=CaO+CO₂ .J. Appl. Chem. Biotechnol. 1973, 23, 733-742) and the zero emission coalalliance process, consisting of hydrogasification of coal fueled by theheat of the carbonation reaction (Tinkler, M. J.; Cheh, C. Towards aCoal-capable Solid Oxide Fuel Cell System. Proceedings of the 26^(th)International Technical Conference on Coal Utilization and Fuel Systems;Clearwater, Fla., Mar. 5-8, 2001; pp 569-570). The gas-solid CaO-CO₂reaction proceeds through two rate-controlling regimes. The first regimeinvolves a rapid, heterogeneous chemical reaction. In the second regime,the reaction slows down due to the formation of an impervious layer ofCaCO₃. This product layer prevents the exposure of unreacted CaO in theparticle core to CO₂ for further carbonation. The kinetics of the secondregime is governed by the diffusion of ions through the CaCO₃ productlayer. The activation energy was estimated to be 21 kcal/mol below 688 Kand 43 kcal/mol above it for the product layer diffusion, based on thecounter migration of CO₃ ²⁻ and O²⁻ ions through the product layer(Bhatia, S. K.; and Perlmutter, D. D. Effect of the product layer on thekinetics of the CO₂-Lime Reaction. AlChE J. 1983, 29(1), 79-86).

The extent of the carbonation reaction reported in many studies has alsoshown considerable variation. Stoichiometrically, 56 g of CaO shouldreact with 44 g of CO₂ to form 100 g of CaCO₃. This translates to about78.6-wt % capacity for CaO. However, the structural limitations preventthe attainment of theoretical conversion. The extent of carbonation wasonly 23-wt % in 30 minutes at 600° C. (Dedman, A. J.; Owen, A. J.Calcium Cyanamide Synthesis, Part 4.—The reaction CaO+CO₂=CaCO₃ . Trans.Faraday Soc. 1962, 58, 2027-2035). A higher surface area CaO sorbentprovided 55-wt % CO₂ sorption (Bhatia, S. K.; and Perlmutter, D. D.Effect of the product layer on the kinetics of the CO₂-Lime Reaction.AlChE J. 1983, 29(1), 79-86). 64-wt % CO₂ sorption was achieved at 1050°C. temperature and 11.74 atm CO₂ pressure in 32 hours (Mess, D.;Sarofim, A. F.; Longwell, J. P. Product Layer Diffusion during theReaction of Calcium Oxide with Carbon Dioxide. Energy and Fuels. 1999,13, 999-1005). However, the extent of carbonation at lowertemperature/pressure conditions that are more characteristic of CO₂containing gaseous mixtures is absent in their work. The limitation intotal conversion stems essentially from the nature of the initial poresize distribution of the CaO sorbent. Microporous sorbents (pore size<2nm) are very susceptible to pore blockage and plugging due to theformation of higher molar volume product (molar volume of CaO: 17cm³/mol; molar volume of CaCO₃: 37 cm³/mol). CaO sorbents obtained fromnaturally occurring precursors are usually microporous in nature. At theend of the kinetically controlled regime, diffusion processes throughthe product layer control the reaction rate. Similar structurallimitations have prevented calcium-based sorbents from attainingtheoretical conversion for the sulfation reaction between CaO and sulfurdioxide (SO₂) as well (Wei, S.-H.; Mahuli, S. K.; Agnihotri, R.; Fan,L.-S. High Surface Area Calcium Carbonate: Pore Structural Propertiesand Sulfation Characteristics. Ind. Eng. Chem. Res. 1997, 36(6),2141-2148). They suggested that a mesoporous structure, which maximizesporosity in the 5-20 nm pore size range, would be less susceptible topore pluggage. This structure would also be able to provide sufficientsurface area to ensure rapid kinetics. Their modified precipitationtechnique resulted in a mesoporous CaCO₃ structure that also had a highBET surface area determined by nitrogen (60 m²/g). A similar approachcould also enhance the reactivity of CaO sorbents towards thecarbonation reaction, which is the focus of this study.

Lastly, it is important that the CaO sorbents maintain their reactivityover many carbonation and calcination cycles. The conversion of CaOdropped from about 73% in the first carbonation cycle to 43% at the endof the 5^(th) cycle at 866° C. (Barker, R. The Reversibility of theReaction CaCO₃=CaO+CO₂ . J. Appl. Chem. Biotechnol. 1973, 23, 733-742 &Barker, R. The Reactivity of Calcium Oxide Towards Carbon Dioxide andits use for Energy Storage. J. Appl. Chem. Biotechnol. 1974, 24,221-227). Barker suggested that the CaCO₃ layer is about 22 nm thick andhis latter work showed repeated 93% conversion over 30 cycles at 629° C.on 10 nm CaO particles. In another study, cyclical studies conducted ata carbonation temperature of 880° C. and calcination at 860° C. led to adrop in conversion from 70% in the first carbonation to 38% in the7^(th) carbonation step (Kato, Y.; Harada, N.; Yoshizawa, Y. Kineticfeasibility of a chemical heat pump for heat utilization from hightemperature processes. Applied Thermal Engineering. 1999, 19, 239-254).The process described here leads to >95% conversion due to theapplication of novel mesoporous CaO sorbents for CO₂ capture andmaintains their reactivity over repeated cycles of carbonation andcalcination.

Part I (CO2/SO2 Combined Reaction Optimization)

Introduction

Carbon dioxide (CO2) accounts for more than half of the enhancedgreenhouse effect, which is responsible for global warming.′ Theatmospheric concentration of CO2 has increased from 280 ppm before theIndustrial Revolution to −365 ppm today. 2′^(2,3) This is mainly due tothe unabated emission of CO2 as a result of increasing consumption offossil fuels such as coal, oil and natural gas. Point sources, such aselectric utility plants that contribute to about one-third of allanthropogenic CO2 emissions⁴, are ideal candidates for implementing CO2reduction practices due to the relatively high concentration andquantity of CO2 emitted compared to smaller, mobile sources. Coalconsumption leads to high CO2 emissions at these large point sources dueto its dominant use in electricity generation (−52%) and higher energyspecific CO2 emission due to its high carbon to hydrogen contentcompared to other fossil fuels (g CO2BTU).^(s) Comprehensive CO2management scenarios involve a three-step process that includesseparation, transportation and safe sequestration of CO2. Economicanalysis has, shown that CO2 separation accounts for 75-85% of theoverall cost associated with carbon sequestration.⁶ Current CO2separation technologies based on absorption, adsorption, membraneseparation, and cryogenic separation necessitate a low temperatureand/or high pressure of flue gas to enhance the CO2 sorption capacity ofthe sorbent/solvent or the diffusion flux of CO2 through the membrane.However, flue gas is typically characterized by sub-atmospheric pressureand high temperature. Metal oxides are capable of reacting with CO2under existing flue gas conditions, thereby reducing downstream processmodifications. We have detailed elsewhere the advantages of a hightemperature reactive separation process based on the carbonation andcalcination reactions (CCR) of CaO to separate CO2 from flue gas.⁷ Thekey advantage offered by this process is the enhanced CO2 sorptioncapacity (35-70 weight %) exhibited by the high reactivity CaO particlesunder existing flue gas conditions over multiple cycles of CCRs.

Extensive screening of metal oxides has identified CaO as a potentialcandidate for the CCR scheme.⁷ The carbonation reaction of CaO has beenstudied for its role in chemical heat pumps^(8,9), energy storagesystems¹⁰, zero emission coal alliance processes″, and in the enhancedproduction of hydrogen from fossil ^(fuels.′) This reaction typicallygoes through a raid kinetic controlled regime, followed by a slowerproduct-layer diffusion controlled regime.′ Naturally occurringprecursors (limestone and dolomite), are unable to achievestoichiometric conversion in any carbonation step due to the predominantmicroporous structure which is susceptible to pore pluggage and poremouth closure. In contrast, mesoporosity, which dominates the porestructure of precipitated calcium carbonate (PCC), synthesized under theinfluence of negatively charged polyacrylate ions yields greater than90% carbonation conversion.⁷″⁴

For the viability of a CCR process, it is imperative that the CaOsorbent maintain high reactivity over multiple cycles. Previous studiesin the literature have reported the performance of numerous CaO sorbentsover multiple cycles. Abanades and co-workers summarized the CCRexperimental data of previous studies on a variety of CaO sorbentsdiffering in their physical properties. They were able to develop asingle correlation between the extent of carbonation as a function ofthe number of CCR cycles. ^(15,16) These sorbents experienced a similarloss in reactivity towards the carbonation reaction regardless ofdifferences in particles size, reactor types, reaction conditions,sorbent characteristics and cycle times. They observed that the highestCO2 sorption capacity retained by the sorbent was 24 wt % after 20cycles.

Sulfur present in coal oxidizes to SO2 during combustion. Calcium basedsorbents are widely used for the control of SO2 emissions. The twoprincipal calcium utilization processes are low temperature wetscrubbing and high temperature furnace sorbent injection (FSI). In wetscrubbing, SO2 capture occurs through ionic reactions in the aqueousphase. In high temperature (>900° C.) FSI systems, calcium oxideprecursors (dolomite, Ca(OH)2 and limestone) and their calcines reactswith SO2 to form CaSO4 via the heterogeneous non-catalytic gas solidreaction. Sulfation under these conditions has been extensively studiedand simulated using various models.″,″,″ PCC achieves a higher extent ofsulfation (−70%) compared to naturally occurring limestone (−30%) atgreater than 900° C. within a residence time of 700 milliseconds.

The flue gas generated by coal combustion typically contains 10-15% CO2,3-4% 02, 5-7% H2O, 500-3000 ppm SO2 and 150-500 ppm NOx in addition totrace quantities of HCl, arsenic, mercury, and selenium. Separation ofCO2 by its absorption in monoethanolamine (MEA) is currently the mostviable option for commercial scale deployment. However, MEA formsthermally stable salts with SO2 and NOx, which do not decompose underthe regeneration conditions employed in the MEA process. It is necessaryto lower SO2 concentration to below 10 ppm to minimize the loss of thecostly solvent. Economic analysis of this process, based on a parasiticconsumption of MEA of 0.5-2 g MEA/kg CO2 separated, show that the costassociated with CO2 separation lies in the $33-73/ton CO2 avoided.²¹ Asimilar hurdle is posed by SO2 for a CaO based CCR process. CaOundergoes sulfation with SO2 forming CaSO4, which cannot be thermallydecomposed back to CaO within the operating temperature range of theproposed CCR process (400-800° C.) as it requires greater than 1100° C.for its decomposition. Exposure of CaO to higher temperatures leads to aloss in surface area and porosity due to excessive sintering, whichdrastically reduces its reactivity. Eventually, the CaSO4 buildup ineach cycle reduces the regenerative capacity of the CaO sorbent oversubsequent cycles ultimately rendering it inactive. However, literatureon the sulfation of CaO in the temperature range where CaCO3 isthermodynamically stable is scant. Sulfation of calcium species in thistemperature range is crucial for the experiments covered in this paperbecause this study aims to investigate the effect of SO2 on thecarbonation of CaO.

The simultaneous hydration, carbonation and sulfation of reagent grade 5micron CaO particles have been previously investigated for an exposuretime of 2 hours in the 170-580° C. temperature range under differentialconditions.²² Low temperatures favor hydration over sulfation andcarbonation. 380° C. marks the termination of hydration and the onset ofcarbonation and sulfation. Carbonation peaked at 520° C. whereassulfation dominated beyond 580° C. Furthermore, the sulfur species inthe form of CaSO3 peaks at 24% at 300° C. and CaSO4 is the only sulfurspecies above 585° C. A high extent of sulfation has also been attainedby 2 mm sized macroporous (>200 run) CaO particles synthesized by aswelling technique involving water-acetic acid mixtures. ^(23.11) Theauthors attributed the high sulfation extent to the increased access ofSO2 to the particle surface due to the macroporosity of the sorbent. Liet al. investigated combined carbonation and sulfation reactions oncommercial grade calcium hydroxide.²⁵ They carried out these reactionsat 425-650° C. by exposing the fines for 2 seconds under entrained flowconditions. The particles were then collected on a hot filter,maintained between 450-510° C. and further exposed to the gas mixturefor 2 hours. Their results indicate an increasing extent of directsulfation of the carbonated product (CaCO3) with higher residencetime.^(2S) The kinetic analysis and modeling of the reaction between SO2and CaCO3 in the temperature range where CaCO3 is thermodynamicallystable, was studied by Snow et al. and Hajaligol et al. ^(26,27) Theyexploited the higher porosity of calcium oxalate derived CaCO3 toachieve about 90% sulfation in the 400-550° C. temperature range. Tullinand Ljungstrom conducted thermogravimetric studies on the simultaneouscarbonation and sulfation of CaO and CaCO3 for a residence time of10-180 minutes at 860° C. The gas mixture consisted of 30-80% CO2_(i)3000 ppm SO2 and 3-4% oxygen ^(28,29) The initial increase in weight ofthe sorbent was predominantly due to the carbonation reaction, whichoccurs to a higher extent than sulfation for the given inlet gasconcentration levels. Further exposure of the sorbent to the reactantgas mixture results in the direct sulfation of the CaCO3 so formed, andleads to a decrease in the overall extent of carbonation and an increasein sulfation. In other experiments, they show that although both CaO andCaCO3 have similar reactivity towards the sulfation reaction, CaSO4formed due to the direct carbonation of CaCO3 is more porous than theCaSO4 product formed by sulfation of CaO.

These literature studies indicate that carbonation occurs at faster ratecompared to sulfation at temperatures around 700° C. due to higherconcentration of CO2. However, SO2 will eventually react directly withCaCO3 leading to the formation of CaSO4. It is thus imperative to obtainexperimental data on combined carbonation and sulfation reactions of CaOover multiple cycles to identify the process conditions under which theextent of carbonation can be maximized in the presence of SO2.Simultaneous high temperature carbonation and sulfation experiments wereperformed in a Thermogravimetric Analyzer (TGA). The study demonstratesthe effect of solid residence time on the overall extent of simultaneouscarbonation and sulfation.

Enhanced Hydrogen Production Integrated with CO₂ Separation in aSingle-Stage Reactor

There has been a global push towards the development of a hydrogeneconomy. The main premise behind this drastic alteration in our energyusage stems from the fact that the use of hydrogen in portable andmobile applications would be the most environmentally beneficial processthat leads only to the emission of water. However, the biggest issuethat needs to be addressed for the success of the hydrogen-based economyinvolves the source of hydrogen itself. While hydrogen may be consideredas the best “carrier” of energy, there is clearly no hydrogen “wells” onearth. The major processes for hydrogen production from fossil fuelsconsist of steam reforming of methane (SMR), coal gasification,catalytic cracking of natural gas, and partial oxidation of heavy oils.Other processes consist of water electrolysis, thermo chemical waterdecomposition, biological processes, etc. (Rosen and Scott, 1998; Rosen,1996). However, water electrolysis is not a very energy efficientprocess.

Water gas, a mixture of CO, CO₂, H₂O and H₂, is formed by thegasification of coal by sub-stoichiometric air and/or steam.Irrespective of the initial concentration of these four gases, thereversible water gas shift (WGS) reaction gets initiated until the exactratio of the concentration of these gases reaches a particularequilibrium constant KWGS that is a function of temperature. The WGSreaction and its equilibrium constant can be written as:

WGS Reaction: CO+H₂O<=>CO₂+H₂ΔH=−40.6 kJ/mol  (1)

WGS Equilibrium Constant:

$\begin{matrix}{K_{WGS} = {\frac{\left\lbrack {CO}_{2} \right\rbrack \left\lbrack H_{2} \right\rbrack}{\lbrack{CO}\rbrack \left\lbrack {H_{2}O} \right\rbrack} = {812.9 - \frac{{6.628\; e} + 5}{T} + \frac{{1.001\; e} + 8}{T^{2}}}}} & (2)\end{matrix}$

where T is in ° C. From equation (2), it can be observed that K_(WGS)reduces with increasing temperature. This means that processes aimed atconverting coal-derived gas to hydrogen at high temperatures arethermodynamically restricted. While catalysts aid in achieving thisequilibrium, they cannot alter the value of K to provide a higherhydrogen yield. An effective technique to shift the reaction to theright for enhanced hydrogen generation has been to remove hydrogen fromthe reaction mixture. This premise has lead to the development ofhydrogen separation membranes. However, membranes cannot completelyremove hydrogen from the mixture. Any remaining hydrogen would diluteCO₂ after its utilization in either a fuel cell or gas turbine.

Another option for driving the WGS reaction forward is to remove CO₂from the reaction mixture by reacting it with CaO. The carbonationreaction can be written as:

Carbonation Reaction: CaO+CO₂→CaCO₃(ΔH=−183 kJ/mol)  (3)

Under the appropriate reaction temperature, CO₂ concentration can belowered down to ppm levels by reaction (3), thereby enabling the maximumproduction of hydrogen from carbon via reaction (1). By conducting thereaction such that CO is the limiting reactant, we can ensure completeutilization of the fuel as well. Besides these advantages, CO₂ issimultaneously removed from the gas mixture in the form of CaCO₃,thereby improving the purity of the hydrogen stream (the othercontaminant being only water). The spent sorbent can then be calcinedseparately to yield pure CO₂ stream, which is then amenable forcompression and liquefaction before its transportation to sequestrationsites. Calcination reaction, reverse of the carbonation reaction can bewritten as:

Calcination Reaction: CaCO₃→CaO+CO₂(ΔH=+183 kJ/mol)  (4)

The resulting CaO sorbent is recycled to capture CO₂ in the next cycle.This cyclical CCR process can be continued so long as the sorbentprovides a satisfactory CO₂ capture.

To obtain high purity H₂, the WGS reaction is generally carried out intwo stages for: (1) high temperature shift (250-500° C.) using ironcatalysts and (2) low temperature shift (210-270° C.) using copper-basedcatalysts (Gerhartz, 1993; Bohlbro, 1969). Copper based catalysts areextremely intolerant to small quantities of sulfur (<0.1 ppm) and hencethe fuel gases need to be desulfurized upstream of the WGS reactor.Besides, to achieve satisfactory carbon monoxide conversion aconsiderable quantity of high-pressure steam is required. For example,to lower the CO content of the typical fuel gas from 45% (inlet) to 3%(outlet) a total steam addition of 1.18 kg/m³ of the gas is required, ata total pressure of 60 bar and 410° C. (Gerhartz, 1993). The steam to COratio at 550° C. can be as high as 50 during a single-stage operation or7.5 for a more expensive dual-stage process to obtain 99.5% pure H2(David, 1980). This is necessary due to the equilibrium limitationinherent in the WGS reaction. From the point of view of H₂ production,even though higher temperatures lead to improved kinetics, WGS has poorequilibrium conditions at the higher temperatures. However, thecontinuous removal of the carbon dioxide product from the reactionchamber will incessantly drive the equilibrium-limited water-gas shiftreaction forward. This will ensure a high yield and purity of hydrogenwith near stoichiometric amounts of steam needed for the reaction.Besides, the reaction can now be carried out at higher temperaturesleading to superior kinetics in the forward direction. Thus the majorequilibrium related drawback in this process could be overcome. Thecontinuous CO₂ removal can be brought about by the carbonation reactionof a metal oxide to give the corresponding metal carbonate. We haveidentified a high reactivity, mesoporous calcium oxide as the potentialsorbent for the in-situ CO₂ capture given by eqn. 3.

The success of this process would effectively bridge coal gasificationto fuel cell usage and chemical synthesis. Other side benefits of thisprocess involve the potential for removal of sulfur and heavy metalssuch as arsenic and selenium from the fuel gas stream.

Recently, Harrison and co-workers reported a single-stepsorption-enhanced process to produce hydrogen from methane(Balasubramanian et al., 1999; Lopez Ortiz and Harrison, 2001). Theyused the traditional concept of SMR with WGS using Ni-based catalyst toproduce hydrogen, coupled with this novel scheme of in-situ continuousCO₂ capture using a calcium-based dolomite sorbent. They obtained highhydrogen yields with 97% purity (dry basis).

However, they reported a low “calcium” conversion in the sorbent ofabout 50% at the beginning of the breakthrough to about 83% at the endof the test. These conversion calculations are based on only the calciumportion of their dolomite sorbent. Their total sorbent conversion willbe much lower than these values as dolomite does not entirely containcalcium based material. In fact, dolomite comprises of nearly 50 wt. %calcium, which participates in the reaction to some extent, and theremaining portion of the sorbent (mainly magnesium oxide) staysunreacted. Further, they attribute the incomplete conversions of thecalcium material to the concept of pore filling and pluggage at thepore-mouths of these sorbent particles by CaCO₃ product layer,preventing the access of CO₂ in the gas to unreacted CaO surface at thepore interiors.

Harrison and co-workers regenerated the dolomite sorbent in streams ofN₂, 4% O₂ in N₂ and pure CO₂. They had to use high regenerationtemperatures of 800-950° C., especially while using pure CO₂. Exposureof the reforming catalyst to an oxidizing atmosphere (viz. O₂/N₂ or CO₂)while regenerating the sorbent used to oxidize the Ni catalysts to NiO.Hence, the catalyst had to be reduced back to Ni before every cycle orthe sorbent-catalyst mixture had to be separated after every run so thatonly the sorbent is subjected to the regeneration conditions. Further,the temperature of operation can be lowered by regeneration in a pure N₂stream. However, it would not solve the problem of CO₂ separation due tothe formation of a CO₂/N₂ gas mixture. Calcination in a pure CO₂ streamwill result in higher operating temperatures due to the thermodynamiclimitations of the calcination reaction in presence of the CO₂ product.Higher temperatures and the presence of CO₂ during calcination wouldcause the sorbent to sinter. This is in agreement with the results ofmultiple carbonation-calcination cycle tests for dolomite by Harrisonand co-workers (Lopez Ortiz and Harrison, 2001) in pure CO₂ stream(800-950° C.). They observed a decrease in “calcium” conversion from 83%in the 1^(st) cycle to about 69% in the 10^(th) cycle itself. However, amesoporous high surface area calcium based sorbent (precipitated calciumcarbonate, PCC) developed at OSU has undergone 100 cycle experiments.The PCC sorbent has shown 85% conversion in the 1^(st) cycle 66.7% inthe 10^(th) cycle and 45.5% in the 100^(th) cycle towards carbonation.These experiments were carried out in a TGA at 700° C. in a 10% CO₂stream in the carbonation cycle and 100% N₂ gas in the calcinationcycle, with 30 minute residence times for each cycle. Therefore thisproject aims testing this PCC based sorbent towards further enhancingthe WGSR and overcoming some of the problems faced by Harrison andco-workers.

SUMMARY OF THE INVENTION

The present invention includes a calcium oxide, its usage for theseparation of CO₂ from multicomponent gas mixtures and the optimumprocess conditions necessary for enhancing the repeatability of theprocess.

A preferred method for separating carbon dioxide from a flow of gascomprising carbon dioxide comprises the steps of: (1) directing the flowof gas to a gas-solid contact reactor, the gas-solid contact reactorcontains at least one sorbent comprising at least one metal oxide; (2)reacting the carbon dioxide with the at least one sorbent so as toremove the carbon dioxide from said flow of gas, thereby converting theat least one sorbent into spent sorbent; (3) calcining the spent sorbentso as to liberate the carbon dioxide from the spent sorbent, therebyregenerating the sorbent; and (4) repeating the aforementioned steps.

Although any metal oxide may be employed, it is preferred that the atleast one metal oxide is selected from the group consisting of: ZnO,MgO, MnO₂, NiO, CuO, PbO, and CaO. Further, it is preferred that thespent sorbent is a metal carbonate.

It is preferred that the sorbent has a sorption capacity of at leastabout 70 grams of carbon dioxide per kilogram of sorbent. However, it iseven more preferred that the sorbent has a sorption capacity of at leastabout 300 grams of carbon dioxide per kilogram of sorbent. Irrespectiveof the sorption capacity of the sorbent, it is preferred that thesorbent has substantially the same sorption capacity after calcining asthe sorbent had prior to adsorbing the carbon dioxide.

Although any calcination method may be employed, it is preferred thatthe calcining is performed under at least partial vacuum. It is alsopreferred that the calcining is performed by steam.

The present invention includes facilities practicing the aforementionedmethod.

A method for separating carbon dioxide from a flow of gas comprisingcarbon dioxide of the present invention comprises the steps of: (1)directing the flow of gas to a first gas-solid contact reactor, thefirst gas-solid contact reactor containing at least one sorbent, thesorbent comprising at least one metal oxide; (2) reacting the carbondioxide in the flow of gas on the sorbent in the first gas-solid contactreactor so as to remove the carbon dioxide from the flow of gas; (3)directing the flow of gas to a second gas-solid contact reactor when thesorbent in the first gas-solid contact reactor is spent thereby formingspent sorbent, the second gas-solid contact reactor containing at leastone sorbent, the sorbent comprising at least one metal oxide; (4)reacting the carbon dioxide in the flow of gas on the sorbent in thesecond gas-solid contact reactor so as to remove the carbon dioxide fromthe flow of gas; (5) calcining the spent sorbent from the firstgas-solid contact reactor so as to generate carbon dioxide and toregenerate the sorbent; (6) directing the flow of gas to the firstgas-solid contact reactor when the sorbent in the second gas-solidcontact reactor is spent, thereby forming spent sorbent; and (7)calcining the spent sorbent from the second gas-solid contact reactor soas to generate carbon dioxide and to regenerate the sorbent.

Although any calcination method may be employed, it is preferred thatthe calcining is performed under at least partial vacuum. It is alsopreferred that the calcining is performed by steam. This applies to bothgas-solid contact reactors.

Although any metal oxide may be utilized, it is preferred that the atleast one metal oxide is selected from the group consisting of: ZnO,MgO, MnO₂, NiO, CuO, PbO, and CaO.

It is preferred that the sorbent has a sorption capacity of at leastabout 70 grams of carbon dioxide per kilogram of sorbent. However, it iseven more preferred that the sorbent has a sorption capacity of at leastabout 300 grams of carbon dioxide per kilogram of sorbent. Irrespectiveof the sorption capacity of the sorbent, it is preferred that thesorbent has substantially the same sorption capacity after calcining asthe sorbent had prior to adsorbing the carbon dioxide.

The present invention also includes facilities practicing theaforementioned method

A method for regenerating a spent sorbent for carbon dioxide of thepresent invention comprises the steps of: (1) providing a spent sorbent,the spent sorbent comprising metal carbonate; and (2) calcining thespent sorbent so as to liberate carbon dioxide gas and so as toregenerate the spent sorbent thereby forming a sorbent comprising ametal oxide.

It is preferred that the spent sorbent is calcium carbonate. It isfurther preferred that the metal oxide is calcium oxide.

It is preferred that the sorbent has substantially the same sorptioncapacity after calcining as the sorbent had prior to adsorbing thecarbon dioxide.

Although any calcination method may be employed, it is preferred thatthe calcining is performed under at least partial vacuum. It is alsopreferred that the calcining is performed by steam. This applies to bothgas-solid contact reactors.

The present invention includes facilities practicing the aforementionedmethod.

A method for producing a sorbent of the present invention comprises thesteps of: (1) obtaining a structurally altered high surface area calciumcarbonate having a surface area of at least 25.0 m²/g, a pore volume ofat least 0.05 cm³/g, and a mesoporous pore size distribution; and (2)calcining the structurally altered high surface area calcium carbonateso as to produce a sorbent having a surface area of less than 22 m²/g, apore volume of at least 0.005 cm³/g, and a mesoporous pore sizedistribution.

Although any calcination method may be employed, it is preferred thatthe calcining is performed under at least partial vacuum. It is alsopreferred that the calcining is performed by steam. This applies to bothgas-solid contact reactors.

The present invention includes sorbents made according to theaforementioned method.

A sorbent according to the present invention comprising calcium oxidehaving a surface area of at least 12.0 m²/g and a pore volume of atleast 0.015 cm³/g, the calcium carbonate sorbent having sorptioncapacity of at least about 70 grams of carbon dioxide per kilogram ofsorbent.

In addition to the novel features and advantages mentioned above, otherobjects and advantages of the present invention will be readily apparentfrom the following descriptions of the drawing(s) and preferredembodiment(s).

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 depicts the Gibbs Free Energy diagram for the carbonationreaction, CaCO₃→CaO+CO₂, as a function of temperature.

FIG. 2 illustrates the performance of calcium oxide for the carbonationreaction.

FIG. 3 compares the XRD diffractograms of CaO derived from variousprecursors.

FIG. 4 is a schematic diagram of a carbonator reactor for the synthesisof precipitated calcium carbonate.

FIG. 5 shows the change in the pH of the slurry as a function of Ca(OH)₂loading. (500 mL water, 0.0575% N40V® dispersant, 4 scfh CO₂).

FIG. 6 depicts the effect of Ca(OH)₂ loading on the morphology ofPrecipitated Calcium Carbonate (PCC) (500 mL water, 0.0575% N40V®dispersant, 4 scfh CO₂).

FIG. 7 compares the pore size distribution of four CaO precursors.

FIG. 8 compares the conversion of four CaO sorbents under pure CO₂ at650° C.

FIG. 9 illustrates the effect of temperature on the carbonation ofPCC-CaO.

FIG. 10 illustrates the carbonation-calcination cycles on Aldrich CaCO₃and PCC at 700° C.

FIG. 11 shows extended carbonation-calcination cycles on precipitatedcalcium carbonate (PCC) powder at 700° C.

FIG. 12 compares the effect of initial surface area of PCC-CaO to itsreactivity towards the carbonation reaction at 700° C.

FIG. 13 depicts the effect of vacuum calcination on the reactivity ofPCC-CaO towards the carbonation reaction at 700° C.

FIG. 14 provides a flow sheet depicting the integration of the currentprocess in the overall coal-gasifier electric production facility.

FIG. 15 a illustrates thermodynamic data for predicting the temperaturezones for hydration and carbonation of CaO.

FIG. 15 b illustrates the typical equilibrium CO₂ partial pressures(PCO₂) as a function of temperature.

FIG. 16 illustrates thermodynamic data for predicting the equilibriumH2S concentration for CaO sulfidation with varying steam concentration(PTotal—1 atm).

FIG. 17 shows a modified reactor set-up with steam generating unit forinvestigating WGS and carbonation reactions.

FIG. 18 illustrates the set-up for combined vacuum/sweep gas calcinationexperiments allowing the use of larger sorbent samples.

FIG. 19 is a pore size distribution of the HTS and LTS obtained from BETanalysis.

FIG. 20 shows the pore size distribution of various calcium oxideprecursors.

FIG. 21 shows the effect of reaction temperature on the CO conversion(0.5 g HTS catalyst, 3% CO, H2O/CO ration=3, total flow=1.5 slpm).

FIG. 22 shows the extent of reaction equilibrium as a function oftemperature for the WGS reaction.

FIG. 23 is a breakthrough curve of CO conversion using a PCC-HTScatalyst system (T=600 C, 3% CO, 9% H2O, Total flow=1.5 slpm).

FIG. 24 is a breakthrough curve of CO conversion using a LH-HTS catalystsystem (T=600 C, 3% CO, 9% H2O, total flow=1.5 slpm).

FIG. 25 provides a comparison of breakthrough curves for PCC-HTS andLH-HTS systems (T=600 C, 3% CO, 9% H2O, Total flow=1.5 slpm).

FIG. 26 depicts a typical steam generation scenario and use.

FIG. 27 depicts one implementation of one embodiment of the presentinvention.

FIG. 28 depicts one implementation of one embodiment of the presentinvention.

FIG. 29 depicts one implementation of one embodiment of the presentinvention.

FIG. 30 depicts one implementation of one embodiment of the presentinvention.

FIG. 31 depicts one implementation of one embodiment of the presentinvention.

FIG. 32 depicts one implementation of one embodiment of the presentinvention.

FIG. 33 depicts one implementation of one embodiment of the presentinvention.

FIG. 34 depicts one implementation of one embodiment of the presentinvention.

FIG. 35 depicts one implementation of one embodiment of the presentinvention.

FIG. 36 depicts one implementation of one embodiment of the presentinvention.

FIG. 37 depicts one implementation of one embodiment of the presentinvention.

FIG. 38 illustrates thermodynamic data for predicting the temperaturezones for sulfation of CaO as well as the direct sulfation of CaCO3.(Sulfation was considered at 1 atm total pressure, 4% 02 and 10% CO2).

FIG. 39 illustrates CO2 capture capacity of various high temperaturesorbents over multiple carbonation-regeneration cycles.

FIG. 40 provides a typical curve for combined carbonation and sulfationof PCC-CaO for 3 cycles at 700° C. for a residence time of 5 minutes(3000 ppm So2, 10% CO2, 4% O2).

FIG. 41 shows the effect of residence time on the extent of carbonation(initial amount of CaO) of PCC-CaO for multiple cycles at 700 C (3000ppm SO2, 10% CO2, 4% O2).

FIG. 42 shows the effect of residence time on the extent of sulfation(initial amount of CaO) of PCC-CaO for multiple cycles at 700 C (3000ppm SO2, 10% CO2, 4% O2).

FIG. 43 shows the effect of residence time on the ratio of carbonationto sulfation of PCC-CaO for multiple cycles at 700 C (3000 ppm SO2, 10%CO2, 4% O2)

FIG. 44 illustrates the extent of carbonation of PCC-CaO for multiplecycles at 700 C (10% CO2, 4% O2).

FIG. 45 illustrates the extent of sulfation of PCC-CaO for multiplecycles at 700 C (10% CO2, 4% O2).

FIG. 46 shows the effect of residence time on the ratio of carbonationto sulfation of PCC-CaO for multiple cycles at 700 C for varying SO2concentrations (3000-100 ppm So2, 10% Co2, 4% O2).

FIG. 47 shows the effect of reaction temperature on the ratio ofcarbonation to sulfation for increasing residence time (10% Co2, 3000ppm SO2).

FIG. 48 illustrates the effect of reaction temperature on the extent ofcarbonation of PCC-CaO for increasing residence time (10% CO2, 3000 ppmSO2).

FIG. 49 illustrates the effect of reaction temperature on the extent ofsulfation of PCC-CaO for increasing residence time (10% CO2, 3000 ppmSO2).

FIG. 50 provides a flow sheet for the integration of the CCR process ina coal fired utility.

FIG. 51 illustrates the equilibrium partial pressure of CO2 as obtainedby thermodynamics (0-1 atm).

FIG. 52 a illustrates a direct fired calcination configuration inaccordance with one embodiment.

FIG. 52 b illustrates an indirect fired calciner configuration inaccordance with one embodiment.

FIG. 53 illustrates a schematic diagram of one embodiment of a rotarycalciner reactor set-up.

FIG. 54 shows the effect of temperature on LC calcination rate (samplesize: 500 mg; T: 700-750 C; Pvac: 25″ Hg; F_(SG(N2))=50 ml/min).

FIG. 55 shows the effect of temperature on the rate of PCC calcination(sample size: 500 mg; T: 700-750 C; Pvac: 25″ Hg; F_(SG(N2))=50 ml/min).

FIG. 56 shows the effect of vacuum on PCC calcination (Sample size: 500mg; T: 750 C; Pvac: 10-25″ Hg; F_(SG(N2))=50 ml/min).

FIG. 57 shows the effect of sweep gas flow (F_(SG)) (Sample size: 10 g,T: 880 C; P_(VAC): 28″ Hg; F_(SG(He))=0-1000 ml/min).

FIG. 58 shows the effect of diluent gas type (He, N2, Ar) (Sample size:10000 mg, T: 800 C; P_(VAC): 28″ Hg; F_(SG(He, N2, Ar))=120 ml/min).

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT(S)

In accordance with the foregoing summary, the following presents adetailed description of the preferred embodiment(s) of the inventionthat are currently considered to be the best mode.

Chemicals, Sorbents and Gases

Naturally occurring limestone (CaCO₃) and hydrated lime (Ca(OH)₂),synthesized from it were obtained from Linwood Mining and Minerals.Dolomite (CaCO₃.MgCO₃) was procured from the National Dolomite Company.The purity of these ores was above 90%. High purity metal oxides such asZnO, MgO, MnO₂, NiO, CuO, PbO, CaO were obtained from Aldrich ChemicalCompany. Precipitated calcium carbonate (PCC) was synthesized fromLinwood hydrate by the procedure described in a following section. N40V®dispersant, a sodium salt of a carboxylic acid, used in the synthesis ofPCC was obtained from Allied Colloid. The synthesis procedure isdescribed in detail in a following section. N₂ and CO₂ used forcalcination and carbonation experiments were 99.999% and 99.9% pure,respectively.

Sorbent Reactivity Testing and Structural Analysis

The reactivity testing of CaO sorbents for carbonation was carried outin a Perkin Elmer Thermogravimetric Analyzer (TGA-7) apparatus. Thebalance can accurately measure up to 1 microgram. A small sample of thesorbent (5-20 mg) is placed in a quartz boat. The weight of the samplewas recorded every second. The structural properties of CaO sorbents andtheir precursors were tested in a NOVA 2200 analyzer (QuantachromeCompany). The BET surface area, pore volume, and pore size distributionwere measured at −196° C. using nitrogen as the adsorbent.

Screening of Metal Oxides

Metal oxides such as ZnO, MgO, CuO, MnO₂, NiO, PbO and CaO that undergothe CCR scheme in the 800-200° C. temperature range were analyzed fortheir reactivity in a TGA. A powdered sample of these oxides was placedin a quartz pan and pure CO₂ was passed over the sample metal oxide. Thetemperature was then slowly raised and the weight of the sample wascontinuously monitored. An increase in the weight of the sample is anindication of the formation of metal carbonate. FIG. 2 providesexperimental data for the carbonation of lime (Ca(OH)₂) under flowingpure CO₂ gas. With an increase in temperature, the weight of the sampleincreases till the temperature reaches about 890° C. Calcination, whichis thermodynamically favored above 890° C. at 1 atm CO₂ partialpressure, causes a rapid decrease in weight until the sorbent convertscompletely to CaO. When the sample is reheated, the weight starts toincrease again and the process is repeated once more. Besides provingthat CaO is a viable candidate, the data also shows recyclability of thesorbent.

XRD Analysis of CaO Obtained from its Precursors:

CaO was identified as a viable candidate for the carbonation-calcinationreactions. However, a variety of precursors can be calcined to obtainthe CaO sorbents necessary for the carbonation reaction. Common andeconomical precursors include calcium carbonate, calcium hydroxide anddolomite. The other important source of CaO is via the calcination ofsynthesized high surface area precipitated calcium carbonate. In orderto compare the crystal structure of the CaO sorbents obtained from thesesources, XRD patterns were obtained on all the CaO sorbents. FIG. 3depicts these diffractograms (a. Calcined Aldrich-CaO; b. Dolomite-CaO;c. Ca(OH)₂—CaO); d. PCC-CaO; e. Limestone-CaO; and f. Aldrich-CaO). Fromthis figure we can conclude that the crystal structure of the CaOsorbents obtained from numerous sources is identical. Only the XRDpattern corresponding to dolomite-derived CaO shows extra peaks due tothe presence of MgO in the calcined dolomite. Based on the similarity inall the CaO structures, it can be assumed that any difference inreactivity of CaO for carbonation is an artifact of the sorbentmorphology and not due to the chemistry of the gas-solid reaction thatoccurs on the CaO surface.

Precipitated Calcium Carbonate (PCC) Synthesis

Structurally altered high surface area CaO precursors were synthesizedbased on the procedure outlined elsewhere (Fan, L.-S.; Ghosh-Dastidar,A.; Mahuli, S.; Calcium Carbonate Sorbent and Methods of Making theSame. U.S. Pat. No. 5,779,464 and Agnihotri, R.; Chauk, S.; Mahuli, S.;Fan, L.-S. Influence of Surface Modifiers on Structure of PrecipitatedCalcium Carbonate. Ind. Eng. Chem. Res. 1999, 38, 2283-2291). Aschematic diagram of the slurry bubble column used for this purpose isshown in FIG. 4. The carbonator 40 consists of a 2″ OD Pyrex tube 40 a.A porous frit 40 d at the bottom, disposed over glass beads 40 f,provides good distribution of CO₂ 40 g through the slurry 40 c. A K-typethermocouple 40 h inserted in the slurry continuously records the slurrytemperature. A pH probe 40 b monitors the pH of the slurry as thereaction medium changes from a basic to an acidic solution as thereaction proceeds. First, 500 ml of distilled water is poured into thecarbonator, followed by the addition of 0.0575 g of N40V®. 12.8 g ofCa(OH)₂ is added to the solution to provide a loading of 2.56% byweight. This corresponds to a concentration of 16-sat (concentration ofCa(OH)₂ is 16 times its saturation solubility limit). The solubility ofCa(OH)₂ (˜0.16 g/100 g water) leads to a pH of 12 at the start of theexperiment. The remaining Ca(OH)₂ remains suspended in the solution. Theratio of N40V® and Ca(OH)₂ loading is chosen to create a surface chargeof zero on the incipiently formed CaCO₃ particles. The flow of CO₂ 40 einto the carbonator is then started and the pH was continuouslymonitored. FIG. 5 shows the change in pH with reaction time as afunction of Ca(OH)₂ loading. CO₂ dissolved in water provides carbonateions that react with Ca⁺⁺ ions to form CaCO₃ according to the reactionbelow:

Ca²⁺+CO₃ ²⁻→CaCO₃  (3)

CaCO₃ has a much lower solubility in water (˜0.0012 g/100 g water)compared to Ca(OH)₂ and thus precipitates out. As the reaction proceeds,Ca²⁺ ions get depleted, but are continuously replenished by thesuspended Ca(OH)₂. Hence the pH remains 12. As the reaction proceeds,Ca(OH)₂ ultimately gets depleted and the concentration of Ca²⁺ ionscannot be maintained at its solubility limit. On the other hand,continued dissolution of CO₂ gas leads to the accumulation of H⁺ ionscausing the solution to become acidic. Eventually, the pH settles atabout 6.0, corresponding to equilibrium solubility of CO₂ in water atambient temperature. This also signals the end of the carbonation of allCa(OH)₂. The slurry is then removed from the precipitator, vacuumfiltered and stored in a vacuum oven at 90-110° C. for 20 hours tocompletely remove the moisture. Higher Ca(OH)₂ loading requires morereaction time as evident from FIG. 5.

Effect of the Ratio of Ca(OH)₂ and Dispersant on PCC Morphology

Precipitated calcium carbonate can be obtained by the reaction betweencarbonate and calcium ions in solution. It is known that the CaCO₃nuclei that precipitate out have positive surface charge on them thatprevent agglomeration (Agnihotri, R.; Chauk, S.; Mahuli, S.; Fan, L.-S.Influence of Surface Modifiers on Structure of Precipitated CalciumCarbonate. Ind. Eng. Chem. Res. 1999, 38, 2283-2291). The resultingstructure is also microporous in nature. However, the structuralproperties of the synthesized PCC can be altered by the use ofnegatively charged dispersants that neutralize the surface charges. Thismakes the ratio between the Ca(OH)₂ loading and the dispersant used verycritical. Besides, the effect of Ca(OH)₂ loading in the slurry wasstudied to enhance the productivity of the precipitation process bysynthesizing more PCC from the same slurry volume. 8-sat, 16-sat and24-sat were used as Ca(OH)₂ loading levels, all other factors remainingconstant. It can be seen from FIG. 6 and Table 1 that at a concentrationof 8-sat, there is proportionally more dispersant in the slurry causingthe incipiently formed CaCO₃ particles to be negatively charged. Thenegative charge prevents the agglomeration of these nuclei eventuallyleading to the formation of microporous PCC as shown in FIG. 6. Itssurface area is also relatively lower. At a Ca(OH)₂ loadingcorresponding to 16-sat, the ratio of N40V® and CaCO₃ is balanced andthe surface charge on the nuclei is zero. This allows optimalassociation of these nuclei leading to a predominantly mesoporousstructure. The SA of PCC under these optimum conditions is also thehighest at 38.3 m²/g. As the loading of Ca(OH)₂ is raised to 24-sat,there is not enough N40V® dispersant to neutralize the surface charge onall the incipiently formed nuclei. There could possibly be somepositively charged particles. This again creates non-optimum conditionsleading to a loss in SA and PV compared to the 16-sat case. Anotherexperiment was conducted to process a 32-sat Ca(OH)₂ slurry keeping theCa(OH)₂ to N40V® ratio constant. The SA/PV of PCC synthesized from a32-sat slurry was 37.07 m²/g and 0.139 cm³/g respectively; lendingsupport to the fact that higher mass of PCC can be synthesized from thesame amount of slurry.

TABLE 1 Morphological properties of PCC as a function of N40V ®: Ca(OH)₂loading ratio (500 ml water, 0.0575% N40V ® dispersant, 4 scfh CO₂).Ca(OH)₂ loading Surface Area Pore Volume weight % (m²/g) (cm³/g) 1.2813.8 0.03 2.56 38.3 0.14 3.84 36.8 0.11

Pore Structure of CaO Sorbents

CaO sorbents were synthesized by calcining various CaO precursors suchas Linwood calcium carbonate (LC), dolomite (DL), Linwood calciumhydroxide (LH), and precipitated calcium carbonate (PCC). Forconvenience, the oxides derived from these sources are termed as LC-CaO,FCD-CaO (for fully calcined dolomite-CaO), LH-CaO, and PCC-CaO,respectively. The procedure involved heating the precursor in flowingnitrogen beyond the calcination temperature (800-950° C.) for an hourfollowed by its storage in a desiccator. Structural properties such assurface area (SA) and pore volume (PV) of these chemicals are listed inTable 2 and their pore size distributions are shown in FIG. 7. The SA ofnaturally occurring minerals, LC and dolomite was very low, 1.06 and1.82 m²/g, respectively. LH was synthesized by first calcining the LCfollowed by its hydration. LH exhibited a considerably higher SA (13.19m²/g) and PV compared to the LC. The SA of PCC (38.3 m²/g), however, wasthe highest among all precursors. From FIG. 5, we can infer that thestructures of LC, DL and LH are predominantly microporous in nature.Most of the porosity lies in pores below 5 nm in diameter. In contrast,the maximum in PV occurs at 15 nm for PCC and most of its PV originatesfrom mesopores in the 5-25 nm range.

TABLE 2 Morphological properties (surface area and pore volume) ofvarious CaO sorbents and their precursors. Surface Area Pore VolumeSorbent Name (m²/g) (cm³/g) LC 1.1 0.003 LC-CaO 17.8 0.078 Dolomite 1.80.004 FCD-CaO 29.8 0.08 LH 13.2 0.0453 LH-CaO 33.3 0.1 PCC 38.3 0.11PCC-CaO 12.8 0.027

Carbonation of CaO Sorbents

The performance of these four CaO sorbents was tested in a TGA. Theexperimental procedure consisted of placing 6-12 mg of the chosen CaOsorbent in a thin layer in a quartz pan to minimize external masstransfer resistances. The sorbent was then heated in flowing nitrogen(5.0 grade, 99.999% pure) to the desired temperature. The representativetemperatures used in these experiments were 550° C., 600° C. and 650° C.Once the desired temperature was reached, the flow was switched to 100%CO₂ stream. The increase in weight with time was recorded and theconversion of CaO to CaCO₃ was calculated from the increase in weight.Only the data obtained at 650° C. is reported here. The performance ofthe four CaO sorbents, LC-CaO, FCD-CaO, LH-CaO and PCC-CaO at 650° C. isdepicted in FIG. 8. Initially, CO₂ diffuses into the pores of the LC-CaOand the reaction takes place on the CaO surface provided by the pores.The figure shows that there is a rapid increase in weight in the first1-2 minutes. The conversion attained in this kinetically controlledregime depends on the initial surface area of the CaO sorbent. LC-CaOand FCD-CaO attained 40-45% conversion, while LH-CaO and PCC-CaOattained about 60% and 54% conversion, respectively, in this regime.After this regime, conversion increases relatively slowly with time. Theincrease in conversion is only about 2-4% in the next hour for LC-CaOand FCD-CaO. This confirms the susceptibility of micropores to porefilling and pore pluggage described earlier due to the formation of ahigher volume product, CaCO₃. The trend is not as dramatic for the caseof LH-CaO because of its relatively higher initial surface area. Theconversion for LH-CaO increases by another 18% in the diffusioncontrolled regime. However, the increase in conversion for PCC-CaO isabout 34-36% more in the second regime. Since the PCC-CaO structure ismesoporous, the formation of CaCO₃ product layer is not able to plug allthe pore mouths. This in turn allows the heterogeneous reaction to occuron a larger CaO surface. Once the kinetically controlled regime is over,diffusion of ions occurs through a larger area, ultimately leading to ahigher conversion of 88-90% for PCC-CaO. FIG. 9 shows the effect oftemperature on the carbonation of PCC-CaO. It can be seen that theextent of conversion in the kinetic regime is different at differenttemperatures. However, unlike LC-CaO, the conversion at any temperaturedoes not seem to taper off and given sufficient time, PCC-CaO is capableof attaining 90% or higher conversion at all of these temperatures.

Cyclic Calcination and Carbonation

One of the possible hurdles in the utilization of metal oxides for thecarbonation and calcination reaction scheme is its vulnerability tosintering due to the thermal cycling imposed by the cyclical nature ofthese reactions. Cyclical studies were carried out to quantify any lossin reactivity of these sorbents upon multiple cycles. The temperaturechosen for cyclical studies was 700° C. This temperature is sufficientto achieve carbonation in the presence of pure CO₂, and also to calcinethe CaCO₃ so formed after the gas is switched from CO₂ to N₂. A varietyof precursors were first calcined in nitrogen at 700° C. The gas wasthen switched to pure CO₂ and the weight gain continuously tracked.After reaching the ultimate conversion, the gas was switched back to N₂.This process was repeated for 2-3 cycles. The data obtained on AldrichCaCO₃ and PCC undergoing this cyclical study is shown in FIG. 10. It canbe seen that the reactivity of Aldrich CaCO₃ exhibited a gradualdecrease even after the first cycle. In contrast, PCC completelyregained its mass after the first calcination and carbonation cycle. At700° C., we can deduce that the conversion is almost complete (>95%).The figure also shows that the reactivity did not decrease in the secondcycle either. Under the reaction conditions chosen, any sintering didnot seem to adversely affect the sorbent morphology. We continued anextended study of eleven calcination and carbonation cycles lasting overthree days on PCC. The data is provided in FIG. 11. It can be seen thatthe sorbent reactivity remained high and if enough reaction time isprovided, the conversion could reach beyond 90% in every cycle. This isa positive result for the structural viability of this sorbent undermultiple cycles.

Effect of Vacuum Calcination

The effect of initial surface area of CaO sorbents was studied. CaOsorbents were synthesized from PCC under different calcinationconditions. The role of surface area on the extent of carbonation isshown in FIG. 12. Different surface area PCC-CaO sorbents weresynthesized by the calcination of PCC at a range of calcinationtemperature to induce varying degrees of sintering. It can be seen thata higher initial surface area (and its associated pore volume) leads tohigher reactivity and conversion. Thus, it is necessary to identifycalcination conditions that optimize the SA/PV and pore sizedistribution of PCC-CaO. It has been suggested in literature that CaOprocured from the calcination of limestone under vacuum has a higherreactivity. It was observed that under air calcination at 650-800° C.,sharp edges of calcite powder were replaced by rounded surfaces and neckareas indicating severe sintering (Beruto, D., and Searcy, A. W.,“Calcium oxides of high reactivity.” Nature, 1976, 263, 221-222). Theresulting CaO structure was highly crystalline as well. In contrast, thesharp edges of calcite were retained in the CaO obtained under vacuum.The CaO however did not possess a high degree of crystallinity. Thelatter also showed high reactivity towards hydration. Vacuum calcinationleads to the formation of a metastable-nanocrystalline calcia structurewhile calcination in helium atmosphere lead to a stable microcrystallinecalcia structure (Dash, S., Kamruddin, M., Ajikumar, P. K., Tyagi, A.K., and Raj, B., “Nanocrystalline and metastable phase formation invacuum thermal decomposition of calcium carbonate.” Thermochimica acta,2000, 363, 129-135). Beruto et al., [1980] estimated the surface areaand pore volume of limestone based CaO to be about 78-89 m²/g and 0.269ml/g respectively.

The effect of vacuum calcination was studied in this process. Thesurface area of Linwood carbonate increased from 17.79 to 21.93 m²/g andpore volume from 0.07815 to 0.1117 ml/g for calcination under nitrogenand under vacuum, respectively. Similar enhancements were observed forPCC based CaO sorbents as well. It has been observed that PCC-CaO issusceptible to high degree of sintering and the surface area of thesorbent falls off rapidly. Calcination in nitrogen resulted in surfaceareas below 13 m²/g repeatedly. However, vacuum calcination lead to asurface area of 19.84 m²/g and 0.04089 ml/g pore volume. The carbonationcharacteristics are shown in FIG. 13.

Vacuum calcination of PCC followed by the carbonation of PCC-CaO wasrepeated over two cycles. PCC was first vacuum calcined to CaO-1 at 750°C. CaO-1 was carbonated to CC-2 at 700° C. followed by its vacuumdecomposition to CaO-2 that is carbonated to CC-3. The values of surfacearea and pore volume of the sorbent at various stages are provided inTable 3 below:

TABLE 3 Structural properties of Calcium based sorbents undergoingvacuum calcination at 750° C. and carbonation at 700° C. Surface AreaPore Volume (m²/g) (cc/g) PCC 38.3 0.1416 CaO-1 12.63 0.02409 CC-2 6.50.0103 CaO-2 15.93 0.04008 CC-3 2.361 0.004483

The data shows that PCC is susceptible to sintering because the CaOobtained in the first cycle has a surface area of only 12.63 m2/gcompared to 38.3 m2/g of PCC. As expected, pore filling leads to a dropin both properties when CaO 1 carbonates. The extent of carbonation wasbeyond 90%. However, it can be seen that the SA of CaO obtained afterthe second vacuum calcination step, CaO 2, is 15.93 m²/g, which ishigher than the SA of CaO 1. The pore volume of CaO 2 is also higherthan that of CaO 1. These results prove that there is no systematicdecline in SA and PV of sorbents with increasing calcination-carbonationcycles and that this combination is capable of providing a sustainedconversion over many cycles.

The article “Carbonation-Calcination Cycle Using High Reactivity CalciumOxide for Carbon Dioxide Separation from Flue Gas” by Himanshu Gupta andLiang-S. Fan, published on the web Jul. 11, 2002 by Ind. Eng. Chem. Res.2002, 41, 4035-4042 is hereby incorporated in its entirety by reference.

Enhanced Hydrogen Production Integrated with CO2 Separation in aSingle-Stage Reactor

A variety of chemical processes known to generate syngas include:

Steam Gasification: C+H₂O→CO+H₂  (X)

Steam Methane Reforming: CH₄+H₂O→CO+3H₂  (X)

Partial oxidation of Hydrocarbon: C_(x)H_(y)+O₂→CO+H₂  (X)

The flow sheet shown in FIG. 14 integrates the Calcium-based reactiveseparation process under development in this project with a coalgasifier based electric power/chemical synthesis process plant 140. Themain coal gasifier 140 a consists of a high pressure and hightemperature unit that allows contact between coal 140 b, steam 140 e andair/pure oxygen 140 y in a variety of schemes. Boiler feed water 140 dis preheated by passing it through gasifier 140 a prior to steam turbine140 c. Waste from the gasifier is collected as slag 140 z. Typical fuelgas compositions from various known coal gasifiers are shown in Table 4.Once the water gas mixture is formed at the exit of the gasifier 140 a,CaO fines are injected 140 f into the gas duct that react with the CO₂present in the gas mixture leading to the formation of solid CaCO₃. Asthe fuel gas flows past the WGS catalyst monoliths 140 g, the WGSreaction is effected forming more CO₂ in the process. The entrained CaOparticles react with the incipiently formed CO₂ gas, thereby allowingfurther catalysis of the WGS reaction to occur. This process can betailored to attain as high a H₂ concentration as possible. At the exitof the WGS reactor, the reacted CaCO₃ particles are captured using ahigh temperature solids separator 140 h (e.g., a candle filter or a hightemperature ESP) and separated fuel gas stream. The spent solids are nowsent to a rotary calciner 140 k to thermally decompose the CaCO₃ 140 jback to CaO 140 f and pure CO₂ 140 m. The high purity CO₂ gas can now beeconomically compressed 140 l, cooled, liquefied and transported for itssafe sequestration 140 m. The rotary calciner allows the calciumparticles to remain segregated, which is crucial in maintaining asorbent structure characterized by a higher porosity. It was previouslyobserved in our studies that heaping of calcium sorbents leads to alower porosity and consequently a lower reactivity over the nextcarbonation cycle. The calcination of the sorbent can also be effectedunder sub-atmospheric conditions that allow the removal of CO₂ as soonas it is formed from the vicinity of the calcining sorbent, therebyaiding further calcination. This vacuum can be created by means ofejector systems that are widely used in maintaining vacuum in largevacuum distillation units (VDU) in the petroleum refining industry. Lockand hopper combinations and appropriate seals ensure that the sorbentcan be effectively separated from the CO₂ stream and re-entrained in thefuel gas duct. The hydrogen enriched fuel gas 140 i can now be used togenerate electric power in a fuel cell 140 n or used to make fuels andchemicals 140 q without any low temperature clean up. The fuel cell mayreceive a supply of air 140 p and discharge steam 140 o. The hydrogenenriched fuel gas may be sent to gas turbine 140 r used to drivegenerator 140 t to produce electricity and air compressor 140 s toproduce a stream of compressed air. The stream of compressed air may besent to air separator 140 x to produce the air/oxygen of 140 y. Thedischarge from gas turbine 140 t may be sent through heat exchanger 140u prior to being discharged at stack 140 v. The absorbed heat may becollected by steam turbine 140 w to produce additional electricity.

Thermodynamic Analysis

Primarily three important gas-solid reactions can occur when calciumoxide (CaO) is exposed to a fuel gas mixture obtained from coalgasification. CaO can undergo hydration, carbonation and sulfidationreactions with H₂O, CO₂ and H₂S, respectively. These can bestoichiometrically represented as:

Hydration: CaO+H₂O→Ca(OH)₂  (5)

Carbonation: CaO+CO₂→CaCO₃  (6)

Sulfidation: CaO+H₂S→CaS+H₂O  (7)

All these reactions are reversible and the extent of each of thesereactions depends on the concentrations of the respective gas speciesand the reaction temperature. Detailed thermodynamic calculations wereperformed to obtain equilibrium curves for the partial pressures of H₂O(PH₂O), CO₂ (PCO₂) and H₂S (PH₂S) as a function of temperature, for thehydration, carbonation, and sulfidation reactions using HSC Chemistry v5.0 (Outokumpu Research Oy, Finland). The equilibrium calculations werebased on the fuel gas compositions that are typical of the differenttypes of coal gasifiers. The details of the fuel gas mixtures areillustrated in Table 4.

TABLE 4 Typical fuel gas compositions obtained from different gasifiers.(Stultz and Kitto, 1992) Moving Moving Bed Fluidized Entrained EntrainedBed, dry slagging Bed Flow, slurry Flow, dry Oxidant air Oxygen OxygenOxygen Oxygen Fuel Sub Bituminous Lignite Bituminous Bitu- Bitu- minousminous Pressure 295 465 145 615 365 (psi) CO 17.4 46 48.2 41 60.3 H₂23.3 26.4 30.6 29.8 30 CO₂ 14.8 2.9 8.2 10.2 1.6 H₂O . . . 16.3 9.1 17.12 N₂ 38.5 2.8 0.7 0.8 4.7 CH₄ + 5.8 4.2 2.8 0.3 . . . HCs H₂S + 0.2 1.10.4 1.1 1.3 COS

The relationship between the reaction temperature and the equilibriumpartial pressures of H₂O and CO₂ for the hydration and carbonationreactions are shown in FIG. 15 (a). For a typical gasifier moisturecomposition ranging from 12-20 atm (PH₂O) the hydration of CaO occursfor all temperatures below 550-575° C., respectively. By operating abovethese temperatures, the CaO-hydration can be prevented. FIG. 15 (b)shows the typical equilibrium CO₂ partial pressures (PCO₂) as a functionof temperature. From the data in Table 4, it can be inferred that thetypical PCO2 in the gasifiers ranges from 0.4-4.3 atm for entrained flow(slurry) and entrained flow (dry) gasifier systems respectively. Theequilibrium temperatures corresponding to those PCO₂ lie in the830-1000° C. range as shown in FIG. 15( b). Thus, by operating belowthese temperatures, we can effect the carbonation of CaO. For thereversible sulfidation of CaO (eqn 7) the thermodynamic calculationsdepend on the concentration of moisture in the system. Hence, FIG. 16depicts the equilibrium H₂S concentrations in ppm for varying moistureconcentrations (PH₂O) and 30 atm total pressure. For a typical operatingtemperature range of 800-1000° C. the equilibrium H₂S concentration isbetween 5700-1700 ppm respectively for 20 atm PH₂O. Consequently, at800° C. we need more than 5700 ppm H₂S for the sulfidation of CaO tooccur. This number changes to 570 ppm for a PH₂O of 2 atm at 800° C.Thus, by changing the moisture/steam concentration in the system we canprevent the sulfidation of CaO from occurring.

Experimental

Sorbent and Catalyst Characterization

The high and low temperature water gas shift (WGS) reaction catalystswere procured from Sud-Chemie Inc., Louisville, Ky. The high temperatureshift (HTS) catalysts comprises of iron (III) oxide supported onchromium oxide. Precipitated calcium carbonate (PCC) was synthesized bybubbling CO2 through a slurry of hydrated lime. The neutralization ofthe positive surface charges on the CaCO₃ nuclei by negatively chargedN40V® molecules forms CaCO₃ particles characterized by a higher surfacearea/pore volume and a predominantly mesoporous structure. Details ofthis synthesis procedure have been reported elsewhere (Agnihotri et al.,1999). Hydrated lime from a naturally occurring limestone (LinwoodHydrate, LH) and a naturally occurring limestone (Linwood Carbonate, LC)was obtained from Linwood Mining and Minerals Co.

The sorbents and catalyst were analyzed to determine their morphologiesusing a BET analyzer. The BET surface areas, pore volumes, and pore sizedistributions of the catalysts and sorbents were measured at −196° C.using nitrogen as the adsorbent in a Nova 2200 Quantachrome BETanalyzer. Special care was taken to ensure that all samples were vacuumdegassed at 250° C. for 5 hours prior to BET analysis.

WGS Reactor Setup

A reactor setup was designed, underwent several iterations and wasassembled to carry out water gas shift reactions in the presence of CaOand catalyst. The reactor design assembly used to carry out theseexperiments is shown in FIG. 17. This setup enables us to carry out boththe water gas shift reaction in the presence of CaO as well as theregeneration of the sorbent in flowing gas such as nitrogen and/orsteam. The setup 170 consists of a tube furnace 170 p, a steel tubereactor 170 a, a steam generating unit 170 c, a set of gas analyzers forthe online monitoring of CO and CO₂ concentrations 170 n, a condenser170 m to remove water from the exit gas stream and a high pressure watersyringe pump 170 b.

All the reactant gases (H₂, CO, CO₂, and N₂) are metered using modifiedvariable area flowmeters 170 e-h respectively. The syringe pump is usedto supply very accurate flow-rates of water into the heated zone of thesteam-generating unit in the 0.01-0.5 ml/min range. Once the steam isgenerated, it is picked up by the CO/N₂ gas mixture 170 i and enters themain reactor where the sorbent/catalyst mixture 170 o is loaded. All thelines connecting the steam-generating unit to the main reactor areheated using heating tapes. The steam generator is also packed withquartz wool 170 d in order to distribute the water drops as they enterinto the heating zone. The packing is utilized in order to providegreater surface for water evaporation and to dampen out fluctuations insteam formation. The main problem with a fluctuating steam supply isthat the gas analyzers used to measure the exit CO and CO₂concentrations are sensitive to gas flow rates. Even though the steam isbeing condensed out before the gas is sent into the analyzers, surges inthe steam supply still affect the overall gas flow rate, causing the COand CO₂ readings to fluctuate. The packing ultimately ensures a morecontinuous and constant overall gas flow rate into the main reactor andinto the analyzers. Thermocouple 170 k is used to monitor thetemperature inside reactor 170 a. Any extra gas inlets of reactor 170 aare blocked 170 l.

A steel tube reactor is used to hold the Ca-based sorbent and catalyst,and is kept heated using a tube furnace. The sorbent loading unit of thereactor is detachable which enables easy removal and loading of thesorbent and therefore minimizes the sorbent loading time between runs.Also, the sorbent can be changed without having to cool down the entirereactor. The gas mixture 170 j entering the reactor is preheated to thereaction temperature before contacting the sorbent/catalyst particles.The gases exiting the reactor first flow through a condenser in order toseparate out the moisture and then to a set of gas analyzers.

Sub Atmospheric Calcination Reactor Setup

Once the Calcium based sorbent has reacted with the CO₂ being producedby the WGSR, the sorbent has to be regenerated for further use insubsequent cycles. During the regeneration of the sorbent, carbondioxide is released from the sorbent. In order to minimize the necessityfor further treatment of this released CO₂ before sending it tosequestration sites, it is necessary to regenerate the sorbent such thata pure stream of CO₂ is released. Vacuum calcination provides one methodfor ensuring that concentrated streams of CO₂ are release in theregeneration phase. The detailed setup is shown in FIG. 18. This setup180 was assembled to handle the regeneration of large quantities ofsorbent (˜10-20 g per batch). The setup includes an alumina tube reactor180 b, which would hold the sorbent samples in a split tube furnace 180c that provides the heat necessary to calcine the sorbent 180 d, two NonDispersive Infra Red (NDIR) analyzers 180 k-l to monitor the CO₂concentration (ranges 0-2500 ppm and 0-20%) and two vacuum pumps 180 fand 180 i. 10 g of sorbent yields about 2.4 L of CO₂ at atmosphericpressure and temperature over the entire decomposition process. This gasneeds to be diluted with air in order to ensure that the CO₂concentration lies in the detection range of the CO₂ analyzers. VacuumPump 180 f is a dry vacuum pump procured from BOC Edwards capable ofachieving vacuum levels as low as 50 mtorr and gas flowrates of 6 m³/hr.The CO2 analyzers have their own inbuilt pumps and are capable ofdrawing up to 2LPM from the header for online CO2 analysis. The secondpump 180 i is a smaller dry pump and is put in place to ensure thatthere is no pressure buildup in the ¼″ lines connecting the vacuum pumpto the analyzers. Pump 180 i discharges to vent 180 j. The temperatureof the furnace is controlled with a thermocouple inserted into thecentral zone of the furnace. The temperature of the reactor was alsomonitored using a second thermocouple inserted into the center of thealumina tube. The setup is also capable of combining vacuum calcinationwith flow of sweep gas 180 a. As it may not be feasible to supply verylow vacuum levels for the calcination of the sorbent in industrialsettings, it may be necessary to study the calcination process incombination with the addition of various sweep gases such as N2/steam.Pressure gauges 180 e, h and volumetric flow meters are included tomonitor the vacuum pressure in the reactor, the pressure in the ¼″ linesand the flows of the sweep gases into the calciner and the flow of theair 180 g used in the dilution of the exhaust CO₂ before sending it tothe analyzers. The analyzers are also connected to a data acquisitionsystem 180 m that can record analyzer readings every second.

Introduction

CaO (s)+CO2 (g)→CaCO3 (s) (Carbonation) ΔH=−178 kJ/mol

CaCO3 (s)→CaO (s)+CO2 (g) (Calcination) ΔH=+178 kJ/mol

Thermodynamics of Calcination

The thermodynamics of calcination, evaluated by HSC chemistry software,is represented in the form of equilibrium partial pressure of CO2 as afunction of temperature (FIG. 1). It indicates that carbonation isfavored under process conditions above the curve and calcination occursat conditions below the equilibrium curve.

It has been amply demonstrated that the ultimate CO2 capture capacity(W) of most sorbents employed at high temperature monotonically fallswith increasing number of CCR cycles (Abanades and Alvarez, 2003, Iyeret al., 2004). While numerous studies have been conducted on thecarbonation reaction to detail the sorbent reactivity, kinetics,mechanism and its mathematical modeling, sufficient emphasis has notbeen placed on the calcination process, as it relates to this CO2separation process. Current calciner designs primarily involve thecombustion of fuel with air inside the rotating tube to supply thesensible heat and heat of calcination directly. The exiting gases, stilldominated by nitrogen, are enriched in CO2_(i) which is released fromthe calcining limestone. However, this design is not amenable togenerating a pure CO2 stream. It is thus imperative that the calcinationdesigns and methods be optimized to maintain the sorbent structure tomaximize reactivity, in a way that the purity of the eventual CO2 streamis not compromised.

Calcination Configurations

The CCR scheme can be carried out in two modes of operation viz.temperature and pressure swing and any combination thereof. Calcinationcan be induced by either increasing the temperature of the carbonatedproduct or by reducing the PCO2 in the calciner such that the processconditions fall below the thermodynamic equilibrium curve. FIG. 2 showsvarious configurations of the calciner operation which detail the modeof heat input to the calciner. The usage of air in FIG. 2( a) representsdirect calcination, which is representative of the commercialcalcination process mentioned earlier. However, a similar reactor designcan be implemented in the CCR scheme if pure oxygen is used in place ofair. The fuel would then form only CO2 and H₂O due to its combustion.The released CO2 from reacted product and the CO2 from the fuelcombustion. can now be further purified by a simple condensation andremoval of steam. Depending on the fuel used for the direct calcination,other trace gases such as SOx and NOx may be emitted, necessitatingfurther control technologies. In particular, make-up calcium would benecessary to replace the sorbent consumed by SO2 in the calciner Oyer etal., 2004)

Alternatively, in the absence of pure oxygen, the heat of calcinationcan be supplied indirectly as shown in FIG. 2( b). The addition of heatwill induce calcination, which leads to CO2 buildup in the calciner.However, the flow of CO2 out of the reactor is possible only if the PCO2becomes greater than I bar. Thermodynamically, _(Pco2) becomes greaterthan 1 bar only above 890° C. It is well known that high temperaturescause sorbent sintering, which reduces its porosity, thereby leading toa drastic reduction in reactivity. Pressure swing mode of operationenables lowering of the calcination temperature to circumvent thesintering problem. A lower PCO2, required by pressure swing operation,is achieved by either dilution of evolved CO2 or by an overall reductionin pressure of the calciner. For example, a reduction in PCO2 below0.0272 bar would lower the calcination temperature to below 700° C.Lowering PCO2 can be accomplished by flowing diluent gas through thecalciner. However, only steam is an acceptable diluent gas since anyother gas such as air, nitrogen, etc. will mix with the evolved CO2defeating the overall objective of isolating a pure CO2 stream. Thereduction in overall calciner pressure, while maintaining 100% pure 002,can be achieved using a vacuum pump which removes CO2 as it evolves fromcalcination.

Literature Review

Past literature studies have shown that CaO resulting from thecalcination under vacuum has a higher reactivity. Beruto and Searcy(1976) observed highly crystalline CaO structure characterized byrounded surfaces and neck areas indicative of severe sintering under aircalcination at 650-800° C. In contrast, the sharp edges of calcite wereretained in the CaO obtained under vacuum. However, this CaO did notpossess a high degree of crystallinity and showed high reactivitytowards hydration. Dash et al., (2000) also observed ametastablenanocrystalline calcia under vacuum as opposed to theformation of stable microcrystalline calcia under helium. Beruto et al.,(1980) estimated the surface area and pore volume of limestone based CaOto be about 78-89 m²/g and 0.269 ml/g respectively. Ingram and Marrier(1963) reported that the rate of reaction varies linearly with thedifference between equilibrium partial pressure and the partial pressureof CO2 surrounding the solid. This further supports the need forsub-atmospheric calcination.

Rao et al. (1989) used thermo-gravimetric reaction data along with agrain model to arrive at the reaction rate constant expression of1.18×10³ exp(−1.85×10/RT). Samtani et al. (2002) investigated thekinetics of calcite decomposition under an atmosphere of N2 anddetermined an activation energy of 192.5kJ/mol and an 1 nA of 20.73(where A is the pre-exponential value of the Arrhenius rate law).Calcite was determined to undergo a zero-order decomposition mechanism,and further investigation into the effect of flow rate, heating rate andsample size did not yield any deviation in the kinetic parameters andmechanism of the process.

Steam enhances calcination due to better thermal properties compared toair and possible catalysis of the reaction. Berger (1927) showed a 30%increase in calcination rate due to steam compared to air in the600-1000° C. Further, the CaO resulting from steam calcination exhibitedhigher rates of slaking, indicating a higher reactivity product comparedthe CaO obtained from air calcination. MacIntire and Stansel (1953)obtained complete calcination at 700° C. as opposed to only 1.6% in air.The catalytic effect of steam, attributed to an activated calciumbicarbonate intermediate species, was responsible for a 20% increase incalcination rate at 834° C. (Terry and McGurk (1994)). Wang and Thompson(1995) hypothesized that the catalysis by steam occurs by the surfaceadsorption of H₂O that weakens the CaO-CO2 bond. Dynamic XRD studiesindicated a 35% increase in conversion due to an addition of 0.77% steamover that in dry helium.

This study demonstrates the role of calcination temperature, level ofvacuum, thermal properties of diluent gas, effect of diluent flow, onthe kinetics of calcination and the morphology of the resultant CaOsorbent. Previous studies have detailed the development of structurallyaltered precipitated calcium carbonate (PCC) sorbent which shows higherreactivity than CaO obtained from the calcination of naturally occurringlimestone (Gupta and Fan, 2002).

Experimental

Reactor Design for Sub-atmospheric Calcination: Rotary Vacuum Calciner

An important objective of this CO2 separation technology is to yield apure/concentrated stream of CO2 during the calcination process. Vacuumcalcination provides one method for meeting this objective. An indirectelectrically heated rotary calciner was designed to carry out thenecessary calcination studies. The calciner design and setup haveundergone a number of modifications however FIG. 3 depicts the finalsetup which incorporates various aspects of prior configurations. Thereactor setup as shown in FIG. 3 was assembled to handle a wide range ofcalcination conditions such as the calcination of 0.5-20 g sorbentsamples under a variety of vacuum, vacuum+sweep gas conditions andcalcination temperatures of up to 950° C. Since early experiments hadshown that sorbent heaping affects both the calcination kinetics as wellas sorbent morphology, the reactor was designed to allow the calcinationto take place under rotary motion which disperses the sorbent therebyminimizing sorbent heaping.

A quartz tube reactor was used to carry out calcination kinetic studies.The reactor tube was designed to have a conical shaped tapered centralzone in order to keep the particles from dispersing axially away fromthe heated zone. Baffles were incorporated to ensure particledispersion. The reactor was placed on two sets of rotary rollers, andwas attached to a motor via a rubber belt mechanism. An electric splittube furnace was used to provide the necessary heat of calcination.Rotary seals enabled the rotation to take place while maintaining thedesired level of vacuum level in the reactor tube. A vacuum level of −28in Hg was achieved in this configuration. The gas exiting the calcinerwas further diluted with air in order to ensure that the CO2concentration fell in the detection range of the CO2 analyzers. A dryvacuum pump (Vacuum Pump 1) procured from BOC Edwards capable ofachieving vacuum levels as low as 50 millitorr and gas flowrates of 6m′/hr was used to supply the necessary reactor vacuum level. The secondpump in the setup is a smaller dry pump and was put in place to ensurethat there was no pressure buildup in the ‘/+” lines connecting thevacuum pump to the analyzers. Two Non Dispersive Infra Red (NDIR)analyzers were used to monitor the CO2 concentration (ranges 0-2500 ppmand 020%). These CO2 analyzers have their own inbuilt pumps and arecapable of drawing up to 2LPM from the header for online CO2 analysis.The temperature of the furnace was controlled with a thermocoupleinserted into the central zone of the furnace. The temperature of thereactor was monitored using a second thermocouple inserted into thecenter of the quartz/alumina tube. Pressure gauges and volumetric flowmeters are included to monitor the vacuum pressure in the reactor, thepressure in the ‘/,” lines and the flows of the sweep gases into thecalciner and the flow of the air used in the dilution of the exhaust CO2before sending it to the analyzers. The analyzers are also connected toa data acquisition system that can record analyzer readings everysecond.

The calcination studies were performed to investigate the role ofcalcination temperature, level of vacuum, thermal properties of sweepgas and effect of gas flow on the kinetics of calcination and themorphology of the resultant CaO sorbent. Experiments were also performedto compare the calcination kinetics of PCC and LC and also toinvestigate the role of sorbent heaping as it relates to scale-upconsiderations of the calciner.

Results and Discussion

Effect of Temperature

Sweep gas is necessary to aid calcination. Prior experiments carried outunder high vacuum conditions in the absence of sweep gas revealed alonger time for calcination. Experiments were performed to determine thepossibility of performing sub-atmospheric calcination in combinationwith the flow of gas. PCC and LC samples of 0.5 g were calcined with asweep N2 gas flow of 50 ml/min under 25″Hg vacuum. FIGS. 4 and 5 showthe conversion plots for LC and PCC at temperature ranges of 700° C. to750° C. The resulting plots show that the calcination time for PCC ismuch lower than that for the naturally occurring LC. At 700° C. PCCtakes about 2000 seconds to fully calcine whereas LC takes −3500seconds. Faster calcination kinetics essentially translates to energysavings in the calcination process which is yet another advantage ofPCC.

Effect of Vacuum Level

Further studies were performed to determine the effect of vacuum on thekinetics of calcination. The kinetic conversion versus time plots curveswere obtained for 0.5 g PCC at vacuum levels of 10″, 15″, and 25″ Hgvacuum with N2 sweep flows of 50 ml/min. The results are plotted in FIG.6 and clearly show that higher vacuum levels translate to fastercalcination times.

Effect of Diluent Flow Rate

It can be observed from FIG. 7 that the calcination time required forlarger samples under pure vacuum conditions is too long. The heat ofcalcination under these conditions is predominantly supplied byradiative means. However, the addition of heat to the calcining sorbentthrough convective means also helps accelerate calcination. This isaccomplished through the use of a preheated diluent gas flow over thecalcining sorbent. FIG. 7 shows the effect of flow rate of diluenthelium gas on the calcination behavior. It indicates that in the absenceof diluent flow, we achieve only 78% calcination in 2000 s. In contrasta steady diluent flow of 120 ml/min attains −93% in 2000 s. As thediluent flow is increased to 500 ml/min, 90% calcination occurs within1200 s. It is also useful to note that an increase in diluent flow to1000 ml/min does not decrease the heat transfer resistancesignificantly.

Effect of Diluent Type

Calcination experiments in the presence of different diluent gases werecarried out to establish the influence of thermal properties (thermalconductivity, heat capacity) of gases on the calcination process. Forexample, at 1000K the thermal conductivity of He (0.354 W/m·K) is higherthan that of N2 (0.0647 W/m·K) (Perry and Chilton, 1997), which couldlead to a difference in the calcination rate FIG. 8 indicates theinfluence of helium, nitrogen and argon on the calcination of 10 gsamples of Linwood carbonate. It can be observed FIG. 8 that heliumindeed causes a faster calcination. In a commercial operation, we cannotuse these particular gases because the gas mixture exiting the calcinerwill consist of CO2, which is evolved from the calcination process andthese diluent gases, thereby defeating the overall purpose of isolatinga pure CO2 stream. However, these experiments lay a foundation for thepotential use of higher thermal conductivity gases such as steam (0.0978W/m·K). Steam has the added advantage of ease of separation fromsteam/CO2 mixtures by its removal by condensation.

Results and Discussions

Catalyst and Sorbent Characterization

The characterization of the high temperature shift (HTS) catalyst in aBET analyzer revealed that the catalyst has a BET surface area of 85m²/g and a total pore volume of about 0.3 cc/g. The majority of thepores were found to occur around 180 Å as evident from the maximum inits pore size distribution plot shown in FIG. 19. In contrast, the lowtemperature shift (LTS) catalyst has a BET surface area of 52 m²/g and atotal pore volume of about 0.1 cc/g. The majority of these pores werefound to occur around 37 Å as evident from the maximum in its pore sizedistribution plot (FIG. 19).

The surface area (SA) and pore volume (PV) of the three different CaOprecursors are provided in Table 5. FIG. 20 shows the pore sizedistribution (PSD) of these precursor fines. It can be seen that LCfines do not have high SA/PV. However, upon calcination and subsequenthydration, the SA/PV of the calcium hydroxide (LH) fines increase as canbe observed for the LH sample. The porosity is maximized in themicroporous range (30-50 Å range). In contrast, the SA/PV of themorphologically altered PCC are much higher. Further, most of theporosity lies in the 100-300 Å range.

TABLE 5 Morphological properties of the natural and synthesized CaOprecursors and the HTS catalyst obtained from BET analyses. Surface PoreVolume Sorbent Area (m2/g) (cc/g) Linwood Carbonate (LC) 1.5 0.004Linwood Hydrate (LH) 13.9 0.050 Precipitated Calcium Carbonate (PCC)49.2 0.170 High Temperature Shift (HTS) catalyst 85 0.3

Water Gas Shift Reaction (WGSR): Catalyst Testing and Analysis

The HTS catalyst was tested for its catalyst activity towards the WGSreaction between 500-700° C. Blank runs (without any sorbent) wereperformed in a reaction mixture comprising of 3% CO and 9% H2O, thebalance being 5.0 grade N2. The total gas flow-rate was maintained atabout 1.5 slpm and the steam/CO ratio was set at ˜3. Typically about 0.5grams of the HTS catalyst was loaded in the reactor prior to each run.The catalyst activity increases monotonically with increasing reactiontemperature. This is evident from FIG. 21 below. The CO conversionincreases from 24.3% at 500° C. to 69.3% at 550° C. It finally reachesabout 80% at 600° C. Beyond 600° C. the conversion does not change muchbut remains steady at ˜78% at 700° C. This might be due to theequilibrium limitations governing the WGS reaction scheme is depicted ineqn (8) below:

CO+H₂O→CO2+H2  (8)

The data were further analyzed to check if the system was operatingwithin the domain of WGS equilibrium. The thermodynamic equilibriumconstant (K) for any temperature for this reaction was computed usingthe software “HSC Chemistry v 5.0” (Outokumpu Research Oy, Finland). Theobserved ratio was computed from the experimental data by obtaining theratio of the partial pressures of the products and the reactants as perthe eqn (9) below:

$\begin{matrix}{{1\text{/}K_{obs}} = \frac{\left( P_{CO} \right)\left( P_{H_{2}O} \right)}{\left( P_{H_{2}} \right)\left( P_{{CO}_{2}} \right)}} & (9)\end{matrix}$

FIG. 22 illustrates the effect of temperature on the ratio of partialpressures (Kobs) obtained from the experimental data. This is comparedwith the thermodynamic equilibrium values (K_(eq)). From the figure itis evident that we are operating in the region that is below thethermodynamic equilibrium. At 500° C. the K_(obs) is 0.028 while thecorresponding K_(eq) is 4.77. K_(eq) monotonically decreases withincreasing temperature. In contrast, K_(obs) increases with temperaturefor our operating conditions. Thus, at 600° C. the K_(obs) increases to1.4 while the K_(eq) moves down to 2.5. This trend continues and it isclearly evident from the figure that the system moves closer toequilibrium as we progressively increase the temperature from 500 to700° C.

Combined Carbonation and Water Gas Shift Reaction:

Sorbent Testing and Analyses

The combined carbonation and WGS reaction was tested in the reactorassembly used for the catalyst testing. The experimental conditions wereexactly identical to that used for testing the catalyst. The runs wereperformed in a reaction mixture comprising of 3% CO and 9% H₂O, thebalance being 5.0 grade N₂. The total gas flow-rate was maintained atabout 1.5 slpm and the steam/CO ratio was set at ˜3. Typically about0.5-1 g of the HTS catalyst was loaded in the reactor prior to each run.Different calcium oxide precursors were tested. Naturally occurringlimestone, Linwood Carbonate (LC) and the corresponding hydrated lime,Linwood Hydroxide (LH) were obtained from Linwood Mining and MineralsCo. The structurally modified calcium carbonate (PCC) was preparedin-house and the details are outlined below.

Sorbent Testing without Catalyst

The sorbents were initially tested for catalytic activity towards WGSRand CO conversion without any HTS catalyst from 500-700° C. This wouldobviate the need for any catalyst in the system. However, detailedinvestigation resulted in very miniscule activity and hence it wasconcluded that HTS catalyst was required for further combined reactiontesting.

Combined Reactions with PCC-HTS Catalyst System

Typically about 0.5 g of HTS catalyst and 1.5 g of PCC were loaded inthe reactor and the temperature was ramped till 700° C. in flowing N₂.This procedure ensured the calcination of the calcium carbonate tocalcium oxide and it was monitored using CO₂ analyzer. Subsequently, thereaction temperature was lowered to 600° C. and the reaction gas mixturewas allowed to flow through the system. The CO analyzer continuouslymonitored the CO flow through the system and the breakthrough curvedepicting the CO conversion with time is as shown in FIG. 23 below. Thesystem gives almost 100% conversion for first 240 seconds (4 min)following which the initial reactivity of the sorbent slowly falls togive about 90% CO conversion at 1000 seconds (16.5 min). The sorbentgradually achieves its maximum loading capacity with time and finally ataround 2500 seconds (42 min) the sorbent reaches its breakthroughloading. Beyond this the CO conversion of 81% corresponds to thatobtained with only the catalyst at 600° C. This can be validated fromFIG. 21.

The system was then switched to pure N₂ flow and the reactiontemperature was increased to 700° C. to drive the calcination of theCaCO₃ formed due to carbonation. Thus the reactions occurring in thesystem are:

Reaction Phase:

WGSR: CO+H₂O→CO₂+H₂  (7)

Carbonation: CaO+CO₂→CaCO₃  (8)

Regeneration Phase:

Calcination: CaCO₃→CaO+CO₂  (9)

The termination of the calcination was ensured by monitoring the CO2released using a CO₂ analyzer. The reaction temperature was againlowered to 600° C. and the sorbent-catalyst system was subjected to thereaction mixture for a second reaction cycle. The 2^(nd) cycle CObreakthrough curve is also depicted in FIG. 23. It is evident from thefigure that the CO conversion is not as superior as in the 1^(st) cycle.The CO conversion monotonically decreases to about 90% in 110 seconds,80% in 240 seconds and gradually to about 50%. It is interesting to notethat at the end of the breakthrough the sorbent-free catalytic COconversion of 81% is not achievable. This could be attributed to theloss in the catalytic activity after the first regeneration cycle. Thisis because the catalyst is subjected to CO₂, an oxidizing atmosphere,during the calcination phase. Thus the deactivated catalyst is not ableto augment the WGS reaction kinetics and hence we see a poor performanceof the sorbent-catalyst system in the 2^(nd) cycle. The solitary sorbenthas been subjected to numerous carbonation calcination cycles and hasshown satisfactory performance Oyer et al, 2004).

Combined Reactions with LH-HTS Catalyst System

Typically about 1 g of the HTS catalyst and 1.3 g of LH were loaded inthe reactor and the temperature was ramped up slowly till 600° C. inflowing N₂. This procedure ensured the calcination of the calciumhydroxide to calcium oxide. Calcium hydroxide decomposes above 400° C.Subsequently, the reaction gas mixture was allowed to flow through thesystem. The CO analyzer continuously monitored the CO flow through thesystem and the breakthrough curve depicting the CO conversion with timeis as shown in FIG. 24 below. The system gives almost 100% conversioninitially to give about 90% CO conversion at 900 seconds (15 min). Thesorbent gradually achieves its maximum loading capacity with time andfinally at around 3000 seconds (50 min) the sorbent has achieved itsbreakthrough loading. Beyond this the CO conversion of 81% correspondsto that obtained with only the catalyst at 600° C. as was shown in FIG.21.

The system was then switched to pure N₂ flow and the reactiontemperature was increased to 700° C. to drive the calcination of theCaCO₃ formed due to carbonation. Subsequently, the reaction temperaturewas lowered to 600° C. and the LH-CaO/catalyst system was subjected tothe reaction mixture for a second reaction cycle. The 2^(nd) cycle CObreakthrough curve is also depicted in FIG. 24. It is evident from thefigure that the CO conversion is not as superior as in the 1^(st) cycle.The CO conversion monotonically decreases to about 90% in 60 seconds,80% in 180 seconds and gradually to about 30%. It is interesting to notethat at the end of the breakthrough the sorbent-free catalytic COconversion of 81% is not achievable. This could be attributed to theloss in the catalytic activity after the first regeneration cycle. Thisis because the catalyst is subjected to CO₂, an oxidizing atmosphere,during the calcination phase. Thus the deactivated catalyst is not ableto augment the WGS reaction kinetics and hence we see a poor performanceof the sorbent-catalyst system in the 2^(nd) cycle. The solitary sorbenthad been subjected to numerous carbonation calcination cycles and hasshown satisfactory performance over few cycles.

Comparison of the PCC and LH Sorbents

FIG. 25 compares the CO conversion breakthrough curves for the PCC andLH sorbent-catalyst systems. The curves are for the 1^(st) reactioncycle. The CO conversion at any given time for PCC-CaO is always higherthan that of LH-CaO. The PCC system gives almost 100% conversion forfirst 240 seconds (4 min) while the LH sorbent system sustains thisconversion only in the initial few seconds. Subsequently, the PCC systemgives about 90% CO conversion at 1000 seconds (16.5 min) followed by 85%in 1600 seconds (27 min). In contrast, the LH system gradually givesabout 90% CO conversion at 900 seconds (15 min) and followed by 85% in1200 seconds (20 min). Both the sorbent systems gradually achieve theirmaximum loading capacity with time and finally at around 2500-3000seconds they reach their breakthrough loading. Beyond this the COconversion of 81% corresponds to that obtained with only the catalyst at600° C. Hence, it is evident from FIG. 24 that the PCC-CaO performancedominates over that of LH-CaO at any given time.

FIG. 26 illustrates the generation 1 MWe of steam.

FIG. 27 illustrates one embodiment of the present invention providing1.002 MWe total capacity.

FIG. 28 illustrates a second embodiment of the present inventionproviding 1 MWe total capacity.

FIG. 29 illustrates another embodiment of the present inventionproviding 1.33 MWe total capacity.

FIG. 30 illustrates yet another embodiment of the present inventionproviding 1.33 MWe total capacity.

FIG. 31 illustrates an alternative embodiment of the present inventionproviding 1.54 MWe total capacity.

FIG. 32 illustrates yet another alternative embodiment of the presentinvention providing 1.07 MWe total capacity.

FIG. 33 illustrates an alternative embodiment of the present inventionproviding 1 MWe total capacity.

FIG. 34 illustrates an alternative embodiment of the present inventionproviding 1 MWe total capacity.

FIG. 35 illustrates yet another embodiment of the present inventionproviding 1.54 MWe total capacity.

FIG. 36 illustrates an alternative embodiment of the present inventionproviding 1 MWe total capacity at 80% CO₂ capture.

FIG. 37 illustrates another embodiment of the present inventionproviding 300 MWe total capacity at 90 CO₂ capture.

CO2/SO2 combined Reaction Optimization

Experimental

Chemicals, Sorbents and Gases

Naturally occurring limestone (CaCO3) was obtained from Linwood Miningand Minerals Company (Linwood Carbonate, LC). Precipitated calciumcarbonate (PCC) was synthesized from Ca(OH)2, obtained from FisherScientific. The pore structure of the synthesized PCC sorbent wastailored using an anionic dispersant, N40V®, obtained from CibaSpecialty Chemicals, Corp. The details of the synthesis procedure isdescribed elsewhere. 7′^(7,14) This structurally modified PCC yields apredominantly mesoporous structure in the 5-20 nm range with a surfacearea (SA) of 49.2 m/g and a pore volume (PV) of 0.17 cc/g obtained byBET analysis. N2 and CO2, obtained from Praxair, Inc were 99.999% and99.9% pure, respectively. Mixtures of O2 and SO2 in N2 were alsosupplied by Praxair, Inc. The BET SA, PV, and pore size distribution(PSD) were measured at −196° C. using nitrogen by a NOVA 2200 analyzer(Quantachrome Company).

Sorbent Reactivity Testing

The reactivity testing of the calcium-based sorbents was carried out ina Thermogravimetric Analyzer (TGA) procured from Perkin ElmerCorporation (Model # TGA-7). A simplified schematic diagram of theexperimental setup is shown elsewhere (Iyer et al., 2004). The domewhich houses the electronic parts of the balance was continuouslyflushed with a pure stream of N2 gas (TGA-N2) to ensure that corrosivegases do not adversely affect the equipment circuitry. The sensitivityof the balance is 1 rig. In these experimental runs, the weight of thesample was recorded every 1-10 second intervals. The gas flows wereaccurately maintained using variable area flow meters, obtained fromCole Parmer Instrument Company. A small sample of the sorbent (about10-12 mg) was placed in a quartz sample holder and brought to 700° C.under nitrogen flow. The temperature of the TGA was then maintained at700° C. throughout the experiment to effect the calcination of PCC.After the calcination step, the valve was switched to allow the flow ofreactant gas mixture over the calcined sorbent (PCC-CaO). An automatedmulti-position valve (VICI Corporation, Model # EMTMA-CE) actuated by aprogrammable electronic timer (VICI corporation, Model # DVSP4) was usedto switch between pure nitrogen stream and the reaction gas mixture atprogrammed time intervals in order to effect the cyclical calcinationand carbonation and sulfation of the sorbent. The alternating flows areadjusted to minimize any variations in weight of the pan/sorbent systemdue to buoyancy changes. The reactant gas mixture enters the TGA fromthe side port and gets diluted by the TGA-N2 stream coming from thebalance dome. The flow of the reactant gas mixture causes an immediateincrease in the weight of the sorbent due to the formation of highermolecular weight products such as CaCO3 and CaSO4. At the end of the setreaction residence time, the automated valve toggles the flow back tothe “calcination nitrogen”. The sorbent weight starts droppingimmediately due to the calcination of the CaCO3 product that is formedin the previous reaction step. The raw data is then analyzed to obtainthe conversion plots.

Results and Discussion

Thermodynamic analysis was carried out to understand the effect ofreaction temperature and gas concentration on the spontaneity of thevarious reactions.

Thermodynamic Analysis

Primarily four gas-solid reactions can occur when calcium oxide isexposed to flue gas from coal combustion. CaO can undergo hydration,carbonation and sulfation reactions with H2O, CO2 and SO2, respectively.In addition, SO2 can react with the CaCO3 formed due to the carbonationreaction, thereby causing direct sulfation of the carbonate. These canbe stoichiometrically represented as:

Hydration: CaO+H₂O→Ca(OH)2  (1)

Carbonation: CaO+CO2→CaCO3  (2)

Sulfation: CaO+SO2+¹/2O2→CaSO4  (3)

Sulfation: CaCO3+SO2+¹/202→CaSO4+CO2  (4)

Thermodynamic calculations were performed to obtain equilibrium curvesfor the partial pressures of H₂O (P_(H2O)), CO2 (P_(CO2)) and SO2(P_(SO2)) as a function of temperature for the hydration, carbonation,sulfation and direct sulfation reactions using HSC Chemistry v 5.0(Outokumpu Research Oy, Finland). The equilibrium curves depicting thetemperature dependent partial pressures of H₂O and CO2 for the hydrationand carbonation reactions are shown in FIG. 1 (a). From theseequilibrium curves, we can predict that moisture does not react with CaObeyond 350° C. in the 5-7% concentration range. At 10% 002, theequilibrium temperature for CaO—CaCO3 system is 760° C. Therefore, thetemperature of the carbonator needs to be kept below 760° C. in order toeffect the carbonation of CaO in a 10% CO2 stream. A temperature of 700°C. offers a reasonable rate of carbonation and calcination reactions andenabled us to carry out multiple CCR cycles under isothermal conditions.Thermodynamic data for the equilibrium temperature versus SO2concentration for the sulfation of CaO and direct sulfation of CaCO3 areshown in FIG. 1( b). The SO2 concentration for the sulfation of CaOsystem is depicted in terms of ppmv for a total system pressure of 1 barat 4% 02. At 700° C., the equilibrium partial pressure of SO2 is 1.84and 5.72 ppt (parts per trillion) for the sulfation of CaO and thedirect sulfation of CaCO3. Since SO2 concentration in the inlet flue gasis in the 500-3000 ppm range, sulfation of CaO and the CaCO3 willdefinitely occur until virtually all SO2 is consumed. Table 1 summarizesthe temperature below which the three reactions are thermodynamicallyfavored at the typical flue gas concentrations at 1 bar total pressure.

Reaction with CaO Hydration Carbonation Sulfation Reactive component ofH2O CO2 SO2 + O2 the flue gas Typical flue gas  5-7%  10-15% 500-3000ppm concentration (vol %) SO2, 3-4% O2 Equilibrium temperature 330-350760-790 1175-1245 below which the reaction can proceed (° C.)

Extended Cyclical Carbonation and Calcination Experiments

Earlier studies from our group have shown that PCC-CaO achieves highconversions (>90%) towards carbonation as compared to 45-60% attained byCaO derived from naturally occurring calcium sources.⁷ Life cycletesting on FCC-CaO, carried out in 100% CO2 for an hour, did not show asignificant drop in reactivity for 2-3 CCR cycles. However, priorliterature indicates a loss in reactivity over a higher number of CCRcycles. We carried out extended isothermal life cycle testing of LC andPCC sorbents at 700° C. The carbonation was carried out in a 10% CO2stream while pure N2 was used for calcination. Each of the CCR steps wasperformed for 30 minutes. The sorption capacity of the sorbent isquantified as wt % CO2 captured by the calcined sorbent. Theoretically,56 grams of unsupported CaO sorbent should react with 44 grams of CO2corresponding to a maximum CO2 sorption capacity of 78.6 wt % at 100%conversion. The wt % capacity of the LC based sorbent towards CO2capture reduces from 58% in the first cycle to 20% at the end of the 50cycle. The microporous structure of LC, being susceptible to porepluggage and pore mouth closure, does not attain high conversion.^(7,24) This is due to the formation of CaCO3, whose molar volume (36.9cc/mol) is higher than that of the reactant CaO (16.9 cc/mol). Incontrast, we see that the conversion of PCC based sorbent over 100cycles is distinctly higher. The capacity of PCC-CaO is 68 wt % in thefirst cycle, which drops to 40 wt % by the 50^(th) cycle and thenslightly to 36 wt % by the 100^(x)′ cycle (6000 minutes on stream). Thehigh reactivity can be attributed to the predominant mesoporousstructure of PCC, which allows the reactant gases to access the entiresurface of particle through the larger pores. The extent of carbonationcontinues to rise significantly beyond the kinetic controlled regime.This fact was ascertained by extending the carbonation reaction time to120 minutes over 40 cycles, during which the sorbent retains 45 wt %capture after 40 cycles (9600 minutes on stream). These results provideevidence that the reactivity of the FCC-CaO is governed solely by thereaction time provided and there is no structural limitation inattaining high conversion.

FIG. 2 depicts graphically the wt % CO2 capture attained by LC, PCC anda host of other high temperature sorbents reported in the literature formultiple CCR cycles.³⁰ While a variety of sorbents have been screenedfor this CCR process, a candidate sorbent that shows consistently highreactivity and sorption capacity over multiple cycles remains to beidentified. The experimental conditions used in the studies referred toin FIG. 5 are detailed in Table 2. This table highlights importantprocess conditions such as carbonation and calcination temperatures,solid residence times, number of cycles, sorption capacities (wt %), andthe CO2 concentration in the gas mixture during the reaction andregeneration steps. FCC-CaO attains 68 wt % increase in 30 minutes and71.5 wt % after 120 min at the end of the first cycle. In contrast,earlier studies have shown a sorption capacity of about 71 wt % (90%conversion) in a pure CO2 stream after 60 min on stream at 650° C.Hence, factors like CO2 concentration, temperature and cycle time play asignificant role in determining the sorption capacity for the samesorbent. The experiments conducted by Barker on 10 micron CaO powderdemonstrate a drop in the sorption capacity from 59 wt % in the firstcarbonation cycle to 8 wt % at the end of 25^(t) cycle.¹⁰ This worksuggests that, due to the formation of a 22 nni thick product layer,particles smaller than 22 nm in diameter should be able to achievestoichiometric conversion. The author later proved this hypothesis byobtaining repeated 93% conversion (73% weight capture) of 10 nm CaOparticles over 30 cycles with a carbonation time of 24 hours under 100%CO2 at 577° C.³¹ In a PbO—CaO based chemical heat rump process, PbOattained 3.6 wt % CO2 capture in the first cycle, decreasing to 1.6 wt %by the 6 cycle and CaO showed a drop in CO2 capture from 53 wt % in the1′⁴ cycle to 27.5 wt % by the 5′^(h) cycle.⁸ A lithium zirconate(Li2ZrO3) based sorbent provided 20 wt % capacity over two cycles.³² Inanother study, researchers at Toshiba Corp. observed that the reactivityof lithium orthosilicate (Li4SiO4) was better than that of lithiumzirconate. Extended cyclical studies performed on lithium orthosilicatesamples attained 26.5 wt % sorption capacity over 50 cycles without anychange in the reactivity.³⁴ Harrison and co-workers developed anenhanced hydrogen production process from the water gas shift reactionby removing CO2 from the gas mixture through the carbonation of CaO. ¹²Dolomitic limestone based CCR process yielded a 35 wt % capacity in thefirst cycle that fell to 11.4 wt % by the 148′^(h) cycle when thecarbonation experiments were performed in pure CO2 at 800° C. andcalcination was conducted at 950° C. An explanation for the drop incapture capacity over multiple CCR cycles has been hypothesized byAbanades and Alvarez based on the changing microporosity within thegrains and the mesoporosity surrounding them due to sintering.¹⁶

Simultaneous Carbonation and Sulfation

The sulfation of CaO and the direct sulfation of the CaCO3 productreduces the CO2 sorption capacity of the CaO due to the formation of“permanent” CaSO4, thereby reducing its efficiency for the CCR processas discussed earlier. This part of the study involves the simultaneouscarbonation and sulfation reactions followed by calcination overmultiple cycles. The goals of this set of experiments are:

(a) to identify the extent of carbonation (X o2) and sulfation (X_(sm))during the simultaneous reactions

(b) to determine and optimize the trend in the ratio of carbonation tosulfation as a function of residence time and reaction temperature

(c) to quantify the reduction in the ultimate carbonation capacity forvarying SO2 concentrations (3000-100 ppm) over multiple CCR cycles

FIG. 3 shows a sample plot of raw data typical for all the experimentsconducted in this section. The x-axis represents the residence time andthe y-axis shows the actual weight of the sorbent at any given instant.In all the experiments, calcination of PCC was carried out typically for20-30 minutes. The residence time for the reaction step was maintainedfor 5 minutes in each of the three cycles for this specific run. Frompoint A to point B, calcination of PCC occurs and the −56% weightremaining confirms that the sorbent at point B is pure CaO. The flowthrough the TGA is then switched to the reactant gas mixture causing theweight of the sorbent to increase from point B to point C due to thecarbonation and sulfation reactions. At the end of the 5-minute reactiontime, gas flow is switched back to N2 to effect the decomposition of theCaCO3 formed due to carbonation in the first reaction cycle causing theweight to drop from point C to point D. In contrast to CaCO3decomposition, CaSO4 formed in the first reaction step remains intact.The extent of carbonation is calculated from the weight loss during thecalcination step C to D, and the extent of sulfation can be estimatedbased on the difference between the weight at point D and that at thestarting point B. Similarly, points D-E-F represent the 2°^(d) cycle andso forth. The trend observed from B to C to D is seen in every cycle ineach experiment. The details of the conversion calculations are reportedelsewhere (Iver et al. 2004).

Carbonation and sulfation occur as heterogeneous non-catalytic gas solidreactions. Higher concentration of CO2 (10% or 100,000 ppm) compared toSO2 (3000 ppm) could result in a higher conversion towards thecarbonation reaction. However, the higher free energy change associatedwith the sulfation reaction thermodynamically favors it over thecarbonation reaction. The process conditions employed can have asignificant impact on the relative rates of these two reactions. Thedata obtained in all experiments has been presented in the form Xco2,X502 and the ratio of carbonation to sulfation, R (Xco2/Xso2) as afunction of reaction residence time

Effect of Residence Time

FIGS. 4-6 show the data obtained on the extent of carbonation andsulfation as a result of simultaneous exposure of PCC-CaO to a gasmixture containing 10% CO2, 3000 ppm SO2 and 4% 02. X-ray diffractionanalysis of the reacted sorbent revealed the presence of CaSO4, CaCO3and CaO only. In the first cycle, XCO2 far exceeds the extent ofsulfation _((Xso2)) during the initial part of the reaction, therebyestablishing the viability of a CCR process for CO2 separation even inthe presence of SO2. XCO2, which increased monotonically in the 0-10minute range, started to fall due to the direct sulfation of the CaCO3formed, consequently leading to a higher _(XS02.) From thermodynamicanalysis presented earlier, it is clear that SO2 concentration greaterthan 5.72 ppt will lead to direct sulfation at 700° C. In fact, _(XS02)attained under simultaneous exposure to CO2 and SO2 is higher than the_(X502) obtained by either the pure sulfation of CaO or the directsulfation of CaCO3 reaction, which are the only possible routes forsulfation. This indicates that the nascent CaCO3, formed due to theparallel carbonation reaction, has a higher reactivity for SO2 than theCaCO3, which forms a part of the stable crystal structure thatcharacterizes the original PCC. After 10 minutes, _(XC02) startsdropping, but it continues to be higher than _(Xs02) until it reaches 40minutes. Beyond 40 minutes, _(XCO2) starts dropping even below XSO2 dueto continued direct sulfation.

FIGS. 4 and 5 depict the effect of residence time over three CCR cycleson _(XCO2) ^(and) Xso2, respectively. PCC-CaO attains a maximum _(XCO2)of −50 wt % at 10 minute residence time. The data in FIGS. 4 and 5 showthat _(XCO2) ^(and) _(XS02) decrease with increasing number of cyclesfor any residence time due to the formation of CaSO4 which reduces theavailability of CaO in the subsequent cycle. The primary reason for thisobservation is the fact that there is a loss in the free CaO due to theformation of non-regenerable CaSO4. FIG. 5 shows that X502 remainsvirtually the same in each of the three cycles until a residence time of10 minutes. In contrast, _(Xco2) shows a significant loss in reactivityover each subsequent cycle in the same duration. For a residence time of60 minutes, XCO2, which was only 22.5% in the first cycle, reduced toalmost zero in the second cycle, indicating a high extent of the directsulfation reaction. In fact, the sorbent is completely spent at the endof the second cycle that it shows no reactivity to either gas in thethird cycle. The overall _(XSO2) for PCC-CaO at the end of three cycleswas 88.2%

FIG. 6 illustrates the ratio R obtained from XCO2 and _(XS02) attainedduring simultaneous carbonation and sulfation. From FIG. 6, we canobserve that the magnitude of R is smaller than that derived from the“individual” reactions and it shifts to 5 minutes instead of 8 minutesseen earlier. This is probably due to the fact that the rate ofsulfation is enhanced due to the simultaneous sulfation of CaO and thehigher reactivity of nascent CaCO3 as explained earlier. From FIG. 6, itis evident that the maximum in the ratio occurs at a reaction time ofabout 5 min for all the three cycles. The magnitude of the ratio fallswith each subsequent cycle and longer residence time. This is due to thedirect sulfation of the calcium carbonate product of carbonationreducing the _(XCO2,) increasing the _(XsO2), and thereby dropping theratio.

Effect of SO2 Concentration

FIGS. 7 and 8 below show the extent of carbonation and sulfationrespectively on PCC-CaO at 700° C. with varying SO₂ concentration from100-3000 ppm over multiple cycles. It is evident from the plots that thecarbonation conversions decrease with increasing cycles and SO2concentrations. The effect of sulfation is very drastic for 3000 ppm andnot so severe with 100-300 ppm range. The extent of sulfation is alsolow in this range as can be observed from FIG. 8. FIG. 9 shows the ratio“R” for increasing CCR cycles with SO2 concentrations varying from 3000to 100 ppm. It is evident from the plots that for each SO2 concentrationcurve there exists a maximum in the ratio, which depends on theresidence time in the system. The ratio is maximum at 160 for a SO2concentration of 100 ppm while it monotonically decreases and reaches ameager value of 5 for 3000 ppm SO2 as seen earlier.

Effect of Temperature

FIGS. 10-12 depict the effect of reaction temperature (500-700° C.) onthe ratio of carbonation to sulfation (R═XCO2/_(XSO2)), the extent ofcarbonation (XCO2) and the extent of sulfation (_(Xso2)) for increasingresidence times (0-30 min). The simultaneous experiments were conductedfor a 3000 ppm SO2, 10% CO2 and 4% 02 stream. As observed in FIG. 10,the ratio (R) decreases with increasing residence time for all reactiontemperatures (except for 700° C.). This is due to the onset of directsulfation of CaCO3 product. It is interesting to note that for any givenresidence time the ratio for 650° C. is the highest, which is followedby 600° C. and subsequently by the values at 700° C. (except at 2 min).The R values at 700° C. are lower than that obtained at temperatures ofboth 600 and 650° C. as sulfation starts to dominate at highertemperatures and the kinetics between these two competing carbonationand sulfation reactions start to play a significant role in determiningtheir ratios. Thus, 650° C. seems to be the optimal temperature tooperate with minimal sulfation effects for a 3000 ppm SO2 and 10% CO2stream. Similarly, the optimal temperature for streams with varying SO2concentrations (3000-100 ppm) needs to be identified.

At 650° C., the ratio starts with a value of 15 for a residence time(RT) of 2 min and subsequently starts to monotonically decrease to about9 for 5 min, 5 for 10 min and finally 2 for 30 min. As illustrated inFIG. 11, the corresponding _(XCO2) is 34% for 2 min, which peaks to 45%at 20 min with a ratio of only 3. It is evident that the extent ofcarbonation is the highest for the temperature of 650° C. for anyresidence time. The only exception is the _(XCO2) of 52% at 700° C. fora 10 min residence time. However the R corresponding to this point isaround 3. Hence, 650° C. is still the preferred temperature of operationwith optimal residence times of 2-5 min giving XCO2 of 34-40% withcorresponding ratios of 15-9 respectively. Thus, working at the optimumtemperature where R as well as _(XCo2) is the highest can maximize theoverall CO2 capture capacity with minimal SO2 effect.

A pilot scale plant, that integrates the CCR process with an actual coalfired combustor will be designed, installed and operated as part of thispilot demonstration. B&W stoker boiler will be used in this process. Aschematic of the process flow diagram is shown in Figure II.C-1. Pleasenote that the process flow diagram could be altered based on futureprocess modifications. It consists of a coal combustion unit thatgenerates actual flue gas. The flue gas is then lowered in its SO2content by the injection of PCC. This FSI mode of sulfur capture hasbeen investigated in an earlier OCDO-OSU sponsored OSCAR project.Results from that project will be factored in to use an optimal Ca/Sratio for SO2 control. The optimum temperature for PCC injection isabout 8001000° C. The entrained fly ash and partially sulfated solids(containing CaSO4 and unreacted CaO) are then physically removed fromthe flue gas through the use of a cyclone. The use of a single cycloneeffectively separates >99% of all solids. This solid mixture is thensafely disposed.

The solids depleted flue gas is then subjected to CO2 removal. This isaccomplished by injecting hot CaO powder through nozzles. The CaOinjection would occur in the 550-700° C. temperature range. Thecarbonator reaction, like sulfation, is extremely fast, with themajority of reactions occurring in less than one second. This translatesto an approximately 1.4 m long flue gas duct requirement under theprocess conditions assuming a residence time of 1 second. The currentequipment under consideration provides ample opportunity for thisresidence time. The carbonation reaction releases tremendous quantity ofheat causing a significant increase in flue gas temperature. However,adequate heat extraction surface would be provided and the actual heattransferred would be accurately quantified. This is one of the mainoutcomes of this project.

The carbonated sorbent is removed from the flue gas by an identicalcyclone, downstream of the carbonation reactor. Attempts would be madeto maintain the flue gas flowing between the two cyclones underisothermal conditions. This will be achieved by providing adequate heattransfer extraction to absorb the heat liberated during carbonation.Hence the flue gas temperature at both cyclones would be the same. Thisis advantageous as identical design can be used for the two cyclones.The use of identical cyclones enhances reliability of the CCR processand reduces design and testing costs associated with cyclones. The fluegas leaving the second “carbonation” cyclone is cooled and sent throughanother lower temperature particle capture device (PCD) and finallyvented out of the building.

The carbonated hot solids are then sent to the calciner, which providesthe heat required to raise the temperature of the solids to 770-830° C.and calcine the carbonated portion of the solids. Prior laboratory dataindicates that the calcination under sub-atmospheric conditions aids inmaintaining a higher porosity CaO sorbent, which also exhibits higherreactivity. A water ejector will generate vacuum for the calcination.The water flowing through the throat of the eductor causes the absolutepressure to fall, allowing the suction of CO2 out of the calciner. Thetwo-phase fluid is then sent to a knockout drum, where water isseparated from the “separated” 002. The water stored in the knockoutdrum is continuously recycled through the water pump for continuousvacuum building.

Conclusions

The enhanced water gas shift reaction for H₂ production with in-situcarbonation was studied at 600° C. using HTS catalyst and calciumsorbents. A naturally occurring calcium precursor (Linwood hydrate, LH)and a modified mesoporous Precipitated Calcium Carbonate (PCC) were usedfor capturing CO₂ for two successive cycles. The PCC system gives almost100% conversion for first 4 min followed by 90% at 16.5 min. Incontrast, the LH sorbent system sustains 100% conversion only in theinitial few seconds and gradually gives about 90% CO conversion at 15min. Experimental evidence clearly shows that the PCC-CaO performancedominates over that of LH-CaO at any given time.

While the invention has been described in connection with what ispresently considered to be the most practical and preferred embodiments,it is to be understood that the invention is not to be limited to thedisclosed embodiment(s), but on the contrary, is intended to covervarious modifications and equivalent arrangements included within thespirit and scope of the appended claims, which are incorporated hereinby reference.

What is claimed is:
 1. A method for separating carbon dioxide and sulfurdioxide from a flow of gas comprising carbon dioxide and sulfur dioxide,said method comprising the steps of: reacting said carbon dioxide andsaid sulfur dioxide with a metal oxide sorbent to convert at least aportion of said metal oxide sorbent to a metal carbonate and at least aportion of said sorbent to a metal sulfate; calcining at least a portionof said metal carbonate to form a reformed metal oxide sorbent andcarbon dioxide; and replenishing said metal oxide sorbent with saidreformed metal oxide sorbent formed during calcination.
 2. The method ofclaim 1, further comprising: separating said metal carbonate from saidgas flow in a cyclone.
 3. The method of claim 1, wherein said conversionof said metal oxide sorbent to metal carbonate and metal sulfate occursat a temperature in the range of about 600° to about 650° Celsius. 4.The method of claim 1, further comprising: removing said carbon dioxideproduced during calcination.
 5. The method of claim 1, wherein saidmetal oxide sorbent has a mesoporous pore size distribution.
 6. Themethod of claim 1, wherein said calcining is conducted under partialpressure.
 7. The method of claim 1, wherein said calcining is performedby steam.
 8. The method of claim 1, wherein said metal oxide sorbent isselected from a group consisting of: ZnO, MgO, MnO2, NiO, CuO, PbO, andCaO.
 9. The method of claim 1, wherein said reformed metal oxide sorbenthas substantially the same sorption as said metal oxide sorbent.
 10. Afacility practicing the method according to claim
 1. 11. A method forseparating carbon dioxide and sulfur dioxide from a flow of gascomprising carbon dioxide and sulfur dioxide, comprising: reacting saidcarbon dioxide and said sulfur dioxide with a sorbent in a gas-solidreactor to convert at least a portion of said sorbent to a metalcarbonate and a portion of said sorbent to a metal sulfate; calciningsaid metal carbonate so as to form a regenerated sorbent; and directingsaid regenerated sorbent to said gas-solid reactor.
 12. The method ofclaim 1, wherein said sorbent is a metal oxide.
 13. The method of claim1, wherein said sorbent is selected from the group consisting of: ZnO,MgO, MnO2, NiO, CuO, PbO, and CaO.
 14. The method of claim 1, furthercomprising: separating said metal carbonate from said flow of gas in acyclone.
 15. The method of claim 1, wherein said regenerated sorbent hasthe same sorption capacity as said sorbent prior to absorbing saidcarbon dioxide.
 16. The method of claim 1, further comprising:sequestering said carbon dioxide formed during calcination.